Water purification

ABSTRACT

The invention provides an efficient method to purify an aqueous solution, typically mine drainage water, especially of anions and cations present in the aqueous solution as dissolved solids, the anions and cations are removed by treatment with a positively charged extractant having at least eight carbon atoms, whereby an unstable emulsion is formed; the unstable emulsion is allowed to break into an extract phase loaded with the anions and cations, and a water phase depleted in anions and cations; a floc inherently forms in the loaded extractant phase and then the loaded extractant phase and floc are separated from the purified water and treated to remove the anions and cations as concentrated useful products; the treated aqueous phase now reduced in anion and/or cation content is also separated from the emulsion as a purified aqueous solution. The extractant phase is preferably recycled. A continuous water purification process is provided.

CROSS-REFERENCE TO RELATED APPLICATIONS

This application is a continuation of U.S. patent application Ser. No.14/082,654, filed Nov. 18, 2013, now U.S. Pat. No. ______, which is acontinuation of U.S. patent application Ser. No. 12/527,325, filed Mar.29, 2010, now U.S. Pat. No. 8,585,903, which was a 371 of PCTApplication No. PCT/US2008/002092, filed Feb. 14, 2008, which directlyclaims the benefit of U.S. Provisional Application Ser. No. 60/901,624,filed Feb. 14, 2007, and is also a continuation-in-part of U.S. patentapplication Ser. No. 11/809,893, filed May 31, 2007, that claims thebenefit of U.S. Provisional Application Ser. No. 60/901,624, filed Feb.14, 2007, Water Purification, Bruce F. Monzyk et al. inventors.

The entire disclosures of these applications are hereby incorporated byreference.

FIELD OF THE INVENTION

The invention provides for large scale water purification that is usefulfor removing anions and cations simultaneously from a large variety ofcontaminated waters.

A particularly useful application is in the purification of acid minedrainage.

BACKGROUND OF THE INVENTION

The demand for decontaminated fresh water has been steadily increasingin the United States and world wide due to increasing populations andincreasing industrialization, mining operations, and agriculture andthis trend is projected to continue. In addition, fresh water sources,such as wells that use aquifer water are increasingly found to becontaminated by these mostly human activities. A particularlyproblematic example is that of mining, especially that of coal mining,where the mining activity has exposed gangue minerals left in the mineand mine tailings to erosion by air, water, and microbial action.Sulfidic minerals, such as pyrite, FeS₂, are commonly found in manygeological strata, and especially in reducing ore bodies such as coalsand metal sulfide.

As described thoroughly in an extensive literature spanning manydecades, coal, metal and other mining operations, and natural weatheringfissures have allowed water, air, and microbial access to these reducingsubstances. These conditions promote the oxidation of the sulfidicminerals to water soluble ferrous sulfate and other metal sulfates;thereby, especially when exposed to air at the surface, producing anacidic discharge or ground water of hundreds and even often above 2000ppm of total dissolved solids (see Tables 2A, 2B, and 2C). Suchcontaminated ground water is unsuitable for most uses, includingmunicipal, industrial, residential, and farming; is foul tasting; isodorous; is toxic to aquatic ecosystems, plants, and animals; and canmobilize additional metal contaminants by acid dissolution of natural orman-made materials.

The problems of water purification and treatments of acid mine drainageand acid rock drainage, collectively referred to as acid mine drainageor AMD water, are well described in the literature. Hereinafter, whenacid mine drainage is discussed, unless mentioned otherwise, the textwill apply to acid rock and natural gas well brines drainage also.Conventional technologies have been found to be unsuitable forprocessing such waters especially where total flow rates exceed 10gal/min and can reach 100,000 gal/min. Dissolved metal cation andcounter anion concentrations in acid mine drainage can be far above thelevels removable by technologies known in the art. Other conventionaltechnologies have major disadvantages including high initial capitalcosts, slow reaction times, high reagent costs, reagent hazards, and/orproduction of waste sludges. But most importantly, they do not removethe major problematic contaminant, sulfate ion. Sulfate ion isresponsible for heavy fouling or scaling during industrial use, foultastes and odors, laxative effects, and very high level of corrosivenessto construction metals and concrete.

BRIEF DESCRIPTION OF THE INVENTION

Broadly, the invention provides for a practical, low-cost, large-scalewater purification process with minimal or no waste generation, suitablefor very high flow rates, even thousands of gallons per minute, providessimultaneous, very rapid anion and cation removal, including sulfateion, which thereby reduces total dissolved solids (TDS), and features asmall equipment foot print, and co-product production. The invention isparticularly suitable for the purification of acid mine drainage or acidrock ground drainage waters (AMD/ARD) without the massive wastegeneration characteristic of prior art technologies. The new process ofthe invention provides recovery of non-toxic, low TDS, and usefulpurified water with metals and sulfate co-products.

This invention provides a means to substantially purify water,especially suited for large volumes of water in a continuous fashion, ina unique manner by the simultaneous removal of anions and cations usinga process based on a particular combination of liquid-liquid extraction(LLX), oil skimming, and flotation (F) technologies. Unique attributesof the process chemistry and associated device design of the invention,referred to as flotation liquid-liquid extraction (F-LLX) (typically thefloc and extractant float), are that it:

-   Provides for the fast removal of these ions at rates 10-100 times    faster than conventional technologies,-   Where only 45 to 90 seconds of contact time is required, and-   Enables the extraction of metal ions normally not extractable or not    well extracted by conventional methods,-   Extracts a broader spectrum of ions, and-   Extracts anions and cations simultaneously (TDS) to lower residual    concentrations than by conventional LLX processing.-   Performs these extractions at mild conditions of pH and temperature    (nominally pH 7-8 and ambient temperature of 1° C. to 99° C.).-   Provides the production of metal ion product concentrates, solid    salts, and/or solids of oxides, hydroxides and/or carbonates.-   Provides the production of sulfate ion product concentrates and/or    solid salts.-   Provides a low cost means to purify large volumes of water    continuously and in high yield (>99%), even thousands of gallons per    minute in a simple manner, and with relatively much smaller    equipment size than conventional treatments.-   Provides a low cost means to purify continuously large volumes of    water that are contaminated with sulfate ion (300-10,000 mg/L SO₄    ²⁻), and even thousands of gallons (0.5-10 Kgal/min) per minute.-   Provides a one step, high yielding, and continuous unit operation    that removes toxic cations, salinity, hardness, TSS, TDS, and/or    acidity, including the very difficult species: ferric, ferrous,    aluminum, nickel, cobalt, manganese, zinc and sulfate ions with acid    neutralization, in a relatively small size of hardware with    concomitant concentration of values to enable their use as products.

For a typical acid mine drainage stream, the process of the inventiondoes not produce large volumes of gypsum/limestone sludge wastes, nordoes it produce membrane or resin back flush wastes, which are producedby conventional technologies. The invention accomplishes wasteprevention by avoiding the addition of voluminous amounts of limestone,lime, slaked lime, dolomite, etc., minerals or wetland muds andvegetation. Instead, an important attribute of the invention is that ithandles the waters for only a few minutes and then releases it; while atthe same time the invention simultaneously concentrates the contaminantsmany fold, normally 10 to 1000 or more times, while continuouslyrecycling the water-immiscible liquid extractant phase. Note that theterm extractant phase and extractant solution are used interchangeablyherein. The continuous regeneration of the extractant phase results inhighly-efficient use of working capital thereby requiring only a smallinventory of the extractant phase. The metal and sulfate productconcentrates are readily processed into useful commodities by well-knownmethods.

The invention includes new compositions of matter consisting offormulations of at least one cationic liquid extractant componentcombined with metal ions to form a colloid consisting of one or moreions selected from oxide, hydroxide carbonate and/or bicarbonate ionscombined with one or more non-basic anions to form an oil soluble phasethat, when combined, form a floc useful for the simultaneous separationand recovery of such solutes. It is especially useful to reduce highmetal ion and sulfate ion concentrations in water, for example as arefound in mining and mineral processing effluent waters, in industrialprocess effluent waters, livestock farm waters, and the like. TDS isreduced from high levels, for example TDS values of 700-5000 mg/L, downto levels that allow water re-use, for example to less than 600 mg/L TDSfor surface water discharge, or less than 250 ppm total dissolved solidsto enable re-use as drinking/potable water, and down to deionized waterfor industrial or non H₂S-form ing potable water use where TDS less thanabout 125 ppm is needed, and with toxic metal ions removed to <1 or even<0.3 ppm. The invention can reduce sulfate ion to <250 or even to <30ppm.

It is believed that the above-mentioned efficient deionization of wateris accomplished by a unique physio-chemical process based on involvingthe fast formation of a new substantially hydrophobic floc materialdescribed above and where such material can be in colloidal form,particulate form, gel form, and/or floc form, and preferably alsocontain oil soluble non-flammable alcohol, ester and/or other oxygenatedhydrocarbon modifier, with a low-viscosity, non-flammable, hydrocarbonliquid diluent, and that is less dense than water and is immiscible inwater so that, left unmixed, it readily separates from water byspontaneous phase disengagement.

Production of waste is minimized in the invention by avoiding the use ofconventional bulking reagents in current practice such as lime, slakedlime, limestone, dolomite, or ferric, ferrous, salts, and/or aluminumsalts or coagulants and/or precipitants, that invariably result inproducing hazardous, unuseable, low-density solid sludge toxic wastesthat lead to dilution of the values.

New products of commerce are also provided that are new composition ofmatter of the formulas:

(R₄N⁺)_(x)(M(II))_(n)(OH⁻)_(z)(SO₄ ²⁻)_(w) (for the case of divalentmetal ions, M(II)  (a)

and

(R₄N⁺)_(x)(M(III))_(y)(OH⁻)_(z)(SO₄ ²⁻)_(w) (for the case of trivalentmetal ions, M(III)  (b)

and consists of a combination of ferric, ferrous, aluminum, and sulfateions in specific range of proportions. This material is formed using theinvention as highly concentrated aqueous solutions ready for use and/oras a crystalline solid material. This new composition is useful forpurification of potable water; purification of agricultural liquidwastes, food wastes, and municipal sewage; purification of industrialwaste water, and the like.

The following disclosure contains a general description where steps ofone or more processes are designated as first, second, third, and so on.This is solely to clearly differentiate process steps that may berepeated in different embodiments from each other.

Broadly the invention includes a method for purifying aqueous solutionsto remove ionic components, consisting of at least one cation and oneanion, comprising:

-   -   a. mixing the aqueous solution with an extractant phase for        typically up to 30 min, preferably 10 min, more preferably 3        min, and most preferably 45-90 seconds to form an unstable        emulsion wherein the extractant phase comprises;    -   (i). A cationic extractant that contains a basic moiety that        forms a neutral to anionic floc with at least one or more of the        cationic components of the aqueous solution and with at least        one anionic component of the aqueous solution, wherein the        extractant comprises a positively charged moiety having at least        8 carbon atoms, preferably 18 carbon atoms, and most preferably        25 carbon atoms, up to about 34 carbon atoms, wherein the carbon        atoms are present as hydrocarbons, and a moiety comprising an        anionic base;    -   (iii). an optional nonflammable hydrocarbon diluent;    -   (iv). an optional nonflammable oxygenated hydrocarbon modifier        for enhancing phase disengagement rate and/or minimizing water        content of the cation and anion loaded extractant portion of the        unstable emulsion;    -   (v) wherein the equilibrium pH of the unstable emulsion is about        2 to about 12 and is controlled by the initial phase ratio of        the extractant and aqueous phases;    -   b. disengaging a first treated aqueous phase and a first loaded        extractant phase from the unstable emulsion, wherein a colloid        and/or floc and/or water immiscible precipitate forms in the        first loaded extractant phase; and    -   c. separating the first loaded extractant phase and floc from        the first treated aqueous phase, wherein the first treated        aqueous phase comprises first purified water.

Typically the positively charged extractant component comprises aquaternary ammonium and/or phosphonium compound selected from the groupconsisting of R₄N⁺, R₄P⁺, and/or an alkylated monoguanadinium compound;where the R groups may differ and are a hydrocarbon consisting of alkylgroups, aryl groups, alkylaryl groups, any combination of these,including atoms of other elements such as N, P, O and S so that thewater solubility is not significantly increased or the monocationiccharge for the whole molecule is not changed, and the charge does notchange with pH up until a pH of about 10, and preferably a pH of about12, and where the minimum carbon number (CN) is >8, preferably >17, andmore preferably >24 up to a total of about 34, and most preferably whereat least one alkyl group in the molecule is branched and/or the wholemolecule has a tripodal structure.

Typically the anionic base is selected from the group consisting of CO₃²⁻, HCO₃ ⁻, or OH⁻, but, depending on the desired products produced andwater contaminants present, also could include PO₄ ³⁻, HPO₄ ²⁻, or H₂PO₄⁻, HS⁻, or S²⁻

The invention provides the unique ability to reduce the total dissolvedsolid levels of waters by simultaneously co-extracting at least one eachof the anionic and cationic components present in the water, and arecaptured from the aqueous solution in the first loaded extractant phaseand floc and the anionic components are one or more of the groupconsisting of sulfate, selenate, nitrate, nitrite, phosphate, arsenate,arsenite, bromate, bromide, perchlorate, iodide, chloride, chromate(VI),permanganate, bisulfide, and sulfide ions, including the protonatedversions of these ions, including those protonated species that wouldrender the ion neutral, and any combination and concentrating level ofthese ions.

Typically the removed cationic component is a metal ion, most oftenselected from one or more of the group consisting of Ni⁺², Fe^(III),Al⁺³, Cu²⁺, Ag⁺, Zn²⁺, Co²⁺, Co^(III), Fe²⁺, Ca²⁺, Mg²⁺, Cd²⁺, Mn²⁺,Pb²⁺, Hg²⁺, Hg₂ ²⁺, CH₃Hg⁺, and Cr^(III), wherein Roman numeralsrepresent variable speciation and the others represent normally aqueousions.

A further embodiment includes neutralization of the acidic component ofthe water simultaneously with the co-extraction of metal ions and anionsinto the extractant phase (E-phase).

A further embodiment includes stripping metal ions and co-extractedanions from the separated first loaded extractant phase, preferably asit exits the extraction decanter, by the steps of:

a. mixing the separated first loaded extractant phase, colloids, andfloc with an aqueous acid to form a second unstable emulsion; whereinmetal ions in the separated first loaded extractant phase and floc arestripped from the emulsion colloids and/or floc of loaded extractantphase making up the unstable emulsion and dissolved into the aqueousacid phase, and

b. disengaging a first loaded aqueous acid phase and a metal ionstripped extractant phase from the second unstable emulsion and wherethe first stripped extractant phase contains at least a portion of theanions extracted into the first loaded extractant phase.

A yet further embodiment includes mixing a metal ion stripped extractantphase with an aqueous solution of anionic base selected from the groupconsisting of CO₃ ²⁻, HCO₃ ⁻, OH⁻, PO₄ ³⁻, HPO₄ ²⁻ or H₂PO₄ ⁻, HS⁻, orS²⁻, wherein HCO₃ ⁻, CO₃ ²⁻ and OH⁻ are most preferred, and where thebicarbonate, carbonate, and/or hydroxide ion loaded extractant phaseproduces a third unstable emulsion, wherein anion, sulfate ion in thecase of AMD feed water, is stripped from the metal ion strippedextractant phase into an aqueous phase and producing a bicarbonate,carbonate and/or hydroxide loaded extractant solution; disengaging theloaded bicarbonate, carbonate and/or hydroxide loaded extractant phaseto yield a regenerated extractant phase stripped of at least a portionof the sulfate content, and normally stripped of more than 90% of itssulfate ion content and preferably stripped of more than 98% of itssulfate ion content, and a third aqueous phase of alkali metal ion,anion solution, alkali, sulfate in the case of AMD water, and/or aslurry of alkali metal ion sulfate and/or alkaline earth metal ionsulfate solution, for the case of alkaline earth Mg, and slurry in thecase of alkaline earth Ca, or a combination of these; and separating theloaded carbonate, bicarbonate and/or hydroxide regenerated extractantphase from the aqueous phase containing the sulfate ion and anyremaining carbonate or hydroxide ions, where preferably most of theloaded basic ion is carbonate where the extraction pH is less than 9 andat least a portion as hydroxide ion where the extraction pH is >9, andwhere the bicarbonate ion component is <10% of the basic anioncomponent.

A yet further embodiment includes a method for treating an aqueoussolution to remove one or more ionic components comprising:

-   -   a. mixing at least a portion of the first treated aqueous phase        from the method above with an extractant phase to form a fourth        unstable emulsion wherein the extractant phase comprises;    -   (i). an extractant that forms a floc with one or more of the        ionic components of the aqueous solution, wherein the extractant        comprises a positively charged moiety having at least 8 carbon        atoms, and a component comprising an anionic base;    -   (iii). an optional medium to low viscosity, nonflammable        diluent;    -   (iv). an optional modifier for increasing phase disengagement        rate and completeness; wherein the equilibrium pH of the        unstable emulsion is about 5 to about 9.    -   b. disengaging a fourth treated aqueous phase and a second        loaded extractant phase from the fourth unstable emulsion,        wherein a floc forms in the second loaded extractant phase; and    -   c. separating the second loaded extractant phase and floc from        the second extracted aqueous phase, wherein the second extracted        aqueous phase comprises second purified water.    -   d. the above steps are repeated twice more to produce a purified        water that has been extracted four times.

Typically this method provides for optional steps where the fourthpurified water is further purified in one or more of an oil/waterseparator, a solid/liquid separator, and/or organic odor sorbent toobtain a further fourth purified water lower in at least one ion fromthe ions still remaining in the first purified water, a purified waterthat is more neutral to slightly alkaline pH, and preferably lower inmore than one ion selected from the list Ni⁺², Cu⁺², Fe^(III), Al⁺³,Zn²⁺, Co²⁺, Co^(III), Fe²⁺, Ca²⁺, Mg²⁺, Cd²⁺, Mn²⁺, Pb²⁺, Hg²⁺, Hg₂ ²⁺,CH₃Hg⁺, Cr^(III), and the like, wherein Roman numerals representvariable speciation and the others represent normally aqueous ions.

The method also provides for removing ionic species wherein the removedionic component is a metal ion selected from one or more of the groupconsisting of cations Ni⁺², Fe^(III), Al⁺³, Zn²⁺, Co²⁺, Co^(III), Fe²⁺,Ca²⁺, Mg²⁺, Cd²⁺, Mn²⁺, Pb²⁺, Hg²⁺, Hg₂ ²⁺, CH₃Hg⁺, Cu²⁺, Cu²⁺,Cr^(III), Au (I and III) and the like, wherein Roman numerals representvariable speciation and the others represent normally aqueous ions, andanions including one or more of sulfate, selenate, nitrate, nitrite,phosphate, arsenate, arsenite, bromate, bromide, perchlorate, iodide,chloride, chromate(VI), permanganate, molybdate, vanadate, and sulfideions, including the protonated versions of these ions, including thoseprotonated species that would render the ion neutral, and anycombination of these ions.

Typically the method provides for capturing at least one ionic componentis captured in the second loaded extractant phase and floc (consistingof colloids, floc, and/or oil dispersible precipitate) and one or moreof the group consisting of cations Ni⁺², Fe^(III), Al⁺³, Zn²⁺, Co²⁺,Co^(III), Fe²⁺, Ca²⁺, Mg²⁺, Cd²⁺, Mn²⁺, Pb²⁺, Hg²⁺, Hg₂ ²⁺, CH₃Hg⁺, Ag⁺,Au(I or III), Cu⁺, Cu⁺, and/or Cr^(III), wherein Roman numeralsrepresent variable speciation and the others represent normally aqueousions, and anions including one or more of sulfate, selenate, nitrate,nitrite, phosphate, organophosphonate, arsenate, arsenite, bromate,bromide, perchlorate, iodide, chloride, chromate(VI), permanganate,molybdate, vanadate, and sulfide ions, including the protonated versionsof these ions, for those ions not forming strong mineral acids,including those protonated species that would render the ion neutral,and any combination of these ions.

Other embodiments provide for a method wherein the first and/or secondloaded extractant phase and floc are further treated to separateresidual water from the loaded extractant phase and floc using adecanter fitted with internal weirs to guide the surface flow ofextractant phase with floc in a narrowing channel such that the flow ofthe floc is maintained as it thickens, and where the flow is maintaineduntil it reaches and flows over an overflow weir designed to promotesuch flow of thickened flocs, most preferably by a about a 15 to about a45° angle, from the vertical, approach ramp, preferably smooth lip andabout a 15 to about 45° exit ramp.

Another embodiment provides for stripping metal ions from at least aportion of the separated second loaded extractant phase by the steps of:

a. mixing the separated second loaded extractant phase and floc with anaqueous acid to form a fifth unstable emulsion; wherein metal ions inthe separated second loaded extractant phase and floc are stripped intothe aqueous acid phase, and

b. disengaging a fifth unstable emulsion with a second aqueous acidphase to form a second metal ion stripped extractant phase separatedfrom the fifth unstable emulsion.

c. optionally recycling the second loaded aqueous strip acid phase tocontact additional volumes of the second loaded extractant phase.

Another embodiment provides for treating the second loaded aqueous acidphase by one or more of an oil/water separator, a solid/liquidseparator, and/or an organic odor sorbent wherein a metal ion saltproduct is obtained.

Another typical method provides for purifying an aqueous solution toremove one or more ionic components, especially residual manganese ionas MnCO₃ particulate, comprising:

-   -   a. mixing at least a portion of the second treated aqueous phase        from the method above with a basic extractant phase to form a        sixth unstable emulsion wherein the extractant phase comprises;    -   (i). an extractant that forms a floc with one or more of the        ionic components of the aqueous solution, wherein the extractant        comprises a positively charged moiety having at least 8 carbon        atoms, and a moiety comprising an anionic base;    -   (iii). an optional diluent;    -   (iv). an optional modifier for modifying phase disengagement;    -   wherein the equilibrium pH of the unstable emulsion is about 8.5        to about 10.5.    -   b. disengaging a third (or more) purified aqueous phase and a        third loaded extractant phase from the seventh unstable        emulsion, wherein a solid suspension and/or floc forms in the        third loaded extractant phase; and    -   c. separating the third loaded extractant phase and floc from        the third purified aqueous phase, wherein the third treated        aqueous phase comprises a third purified water.

A yet further method includes purifying an aqueous solution at higher pHto remove one or more ionic components, especially magnesium ion asMg(OH)₂ particulate, and any residual amount of sulfate ion, comprising:

-   -   a. mixing at least a portion of the third treated aqueous phase        from the above with a base-loaded extractant phase to form an        eighth unstable emulsion wherein the extractant phase comprises;    -   (i). an extractant that forms a floc or particulate with one or        more of the ionic components of the aqueous solution, wherein        the extractant comprises a positively charged moiety having at        least 8 carbon atoms, and a moiety comprising an anionic base,        most preferably hydroxide ion;    -   (iii). an optional diluent;    -   (iv). an optional modifier for modifying phase disengagement;    -   wherein the equilibrium pH of the unstable emulsion is about        10.5 to about 12.    -   b. disengaging an fourth treated aqueous phase and a fourth        loaded extractant phase from the eighth unstable emulsion,        wherein a floc and/or particulate slurry forms in the fourth        loaded extractant phase; and    -   c. separating the fourth loaded extractant phase and floc and/or        particulate from the fourth purified aqueous phase, wherein the        fourth purified aqueous phase comprises the eighth purified        water.

Typically the method provides for treatment wherein the fourth purifiedwater is further treated in one or more of an oil/water separator, asolid/liquid separator, to separate out the Mg(OH)₃ particulate product,to obtain a further fourth purified water low in most ions, includingcations Ni⁺², Fe^(III), Al⁺³, Zn²⁺, Co²⁺, Co^(III), Fe²⁺, Ca²⁺, Mg²⁺,Cd²⁺, Mn²⁺, Pb²⁺, Hg²⁺, CH₃Hg⁺, and/or Cr^(III), wherein Roman numeralsrepresent variable speciation and the others represent normally aqueousions; and anions including one or more of sulfate, selenate, nitrate,nitrite, phosphate, arsenate, arsenite, bromate, bromide, perchlorate,iodide, chloride, chromate(VI), permanganate, and sulfide ions,including the protonated versions of these ions, including thoseprotonated species that would render the ion neutral, and anycombination of these ions, and a ninth separated extractant phase thatis optionally sent to the mixing step of a previous step to load evenhigher concentrations of cation and anion values.

Another embodiment of the invention provides a method for treating anaqueous solution containing ionic components comprising:

-   -   a. mixing the aqueous solution with a first extractant phase to        form a first unstable emulsion wherein the first extractant        phase comprises;    -   (i). one or more of a quaternary ammonium compound comprising        R₄N⁺, an alkylated guanidinium compound, or a quaternary        phosphonium compound;    -   (ii) a carbonate and/or hydroxide component;    -   (iii). an optional diluent; and    -   (iv). an optional modifier for helping the phases disengage;    -   wherein the equilibrium pH of the unstable emulsion is about 2        to about 12.    -   b. disengaging a first treated aqueous phase and a first loaded        extractant phase from the unstable emulsion, wherein a floc        forms in the first loaded extractant phase; and    -   c. separating the first loaded extractant phase and floc from        the first treated aqueous phase, wherein the first treated        aqueous phase is purified water.

A further embodiment provides for a floc composition comprising: an oilsoluble cation, a metal ion, a hydroxide and or oxide, and an anionaccording to the formula,

{(R₄N⁺)_(x)(M(II))_(n)(OH⁻)_(z)(SO₄ ²⁻)_(w)}_(m) (for the case ofdivalent metal ions)

or

{(R₄N⁺)_(x)(M(III))_(y)(OH⁻)_(z)(SO₄ ²⁻)_(w)}_(q) (for the case oftrivalent metal ions), and

mixtures of compositions of the formuli;wherein x=2-4, n=0-1, Y=1-2, Z=2, and w=1-3, and wherein M(III) can beFe(III) or Al, and M(II) can be Fe(II), Ni, Co(II), Cu, Zn, Pb, Cd, orMn;and wherein m and q can be 1 to about 100,000.

An additional floc composition includes a floc composition comprising:an oil soluble cation, a metal ion, a hydroxide and or oxide, and ananion of the formula,

(R₄N⁺)_(x)(M(II))_(n)(OH⁻)_(z)(SO₄ ²⁻)_(w) (for the case of divalentmetal ions M(II)

or

(R₄N⁺)_(x)(M(III))_(y)(OH⁻)_(z)(SO₄ ²⁻)_(w) (for the case of trivalentmetal ions M(III))

and mixtures of materials according to the two formuli;wherein x=2-4, n=0-1, Y=1-2, Z=2, and w=1-3, and wherein M(III) can beFe(III) or Al, and M(II) can be Fe(II), Ni, Co(II), Cu, Zn, Pb, Cd, orMn, and where the material is dispersed in a non water soluble liquidsuch that the floc represents 1 to 100% by weight of the slurry.

A yet further embodiment includes a mixture of metal sulfates of the thecomposition Fe^(III) _(x)Fe^(II) _(y) Al_(z)(SO₄)_(w) where x=0.1, y=0.8and z=0.1, giving a value for w=(0.3+1.6+0.3)/2=1.1, or Fe^(III)_(0.3)Fe^(II) _(1.6)Al_(0.3)(SO₄)_(1.1) for a formula weight of 159g/mole; and having the range of ratios: Fe^(III) _(0.03)Fe^(II)_(0.95)Al_(0.03)(SO₄)_(1.04) (FW of 155 g/mole), to Fe^(III)_(0.95)Fe^(II) _(0.03)Al_(0.5)(SO₄)_(2.2) (FW of 280 g/mole). Typicallythe material, Fe^(III) _(x)Fe^(II) _(y) Al_(z)(SO₄)_(w), produced by theprocess of the invention from acid mine drainage feed water is in anaqueous solution and has a dry weight of at least 0.1%, and preferably 1to 5%, and most preferably 5 to 20%.

An additional embodiment of the invention includes an apparatus forseparating floc consisting of a. a decanter or extractor compartmenthaving an inlet for a mixture at one end of the compartment and anoutlet for aqueous spaced apart from the inlet; and b. a floc weirspaced apart from the inlet, the floc weir comprising a tapered entranceramp, a rounded lip adjacent to the tapered entrance ramp and a flocweir exit slide adjacent to and spaced beyond the rounded lip, whereinfloc flows out of the compartment over the entrance ramp, rounded lipand floc weir exit slide. Typically the entrance ramp is tapered at 15to 45 degrees from the vertical, and the exit slide is tapered at 15 to45 degrees from the vertical. A curved drip point at the lowest point ofthe exit slide is typically used to provide for smooth separation of thefloc from the slide. A rounded lip located between the entrance ramp andexit slide of the floc weir provides a smooth transition for floc flowover this portion of the floc weir, and is typically the highest pointon the floc weir.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1A is a schematic drawing illustrating one part of the overall flowof materials of one aspect of the invention.

FIG. 1B is a schematic drawing illustrating a second part of the overallflow of materials one aspect of the invention. FIG. 1A and FIG. 1B takentogether are intended to form one large figure.

FIG. 1C is a schematic showing the general process steps involved in abroad embodiment of the invention.

FIG. 2 is a schematic illustrating the process for one embodiment of theinvention for the separation of trivalent and divalent metal ions andsulfate ions, simultaneously from contaminated water. Also see the otherfigures and examples illustrating more detailed aspects of the process.

FIG. 3 is schematic diagram showing one typical LLX process flow diagramof the invention illustrating a most preferred apparatus and process fortreatment of acid mine drainage. FIG. 3 also illustrates theconfiguration used in Example 5 (Run #7).

FIG. 4 is a schematic illustrating various aspects of a typicalflotation weir device according to one aspect of the invention.

FIG. 5 is a schematic of a flow guide that guides and facilitates theflow of floc and extractant over a weir including details of its use inthe invention.

FIG. 5A is a schematic drawing showing a side cutaway view of Settler526 with a plate for a flow guide.

FIG. 5B is a schematic drawing showing a side cutaway of view of Settler526 with a block for a flow guide.

FIG. 5C is a schematic drawing of a mixer according to the embodiment ofFIG. 18.

FIG. 6 is a schematic illustrating a typical underflow weir forextractant introduction.

FIG. 7 is a schematic illustrating a typical design having extractantsolution introduction by tubing that is inside the mixing chamber.

FIG. 8 is a schematic drawing illustrating a typical design for anextractant storage chamber and the path of introducing extractant intothe mixing.

FIG. 9 is a graph showing sulfate loading stages determination fromMcCabe-Thiele plots with 5 (v/v) % extractant.

FIG. 10 is similar to FIG. 9 except that sulfate loading stagesdeterminations from McCabe-Thiele plots with 5 (v/v) % extractant areshown.

FIG. 11 is a graph showing other conditions for sulfate loading stagesdetermination from McCabe-Thiele plots with 5 (v/v) % extractant.

FIG. 12 is similar to FIG. 11 except that sulfate loading stagesdeterminations from McCabe-Thiele plots with 5 (v/v) % extractant areshown.

FIG. 13 is a schematic diagram illustrating a sulfate circuit processflow scheme according to one aspect of the invention.

FIG. 14 is a schematic diagram illustrating acid mine drainage processflow scheme 1.

FIG. 15 is a schematic diagram illustrating an acid mine drainageprocess flow scheme 2.

FIG. 16 is a plot of the correlation of Aliquat™ and isodecanolconcentration and the impact of the formulation to post phase separationwater entrainment.

FIG. 17 is a schematic drawing illustrating a presently preferredconfiguration.

DETAILED DESCRIPTION OF THE INVENTION BEST MODE

Broadly, the invention provides for a method for purifying an aqueoussolution by simultaneously removing both cationic and anionic componentswhile neutralizing acidity and lowering total dissolved solids by thesteps of mixing the aqueous solution with a water immiscible liquidextractant phase (E-phase) containing a basic moiety to form anunstable, high surface area emulsion during which time these componentstransfer from the aqueous phase to the E-phase, and acidity in theaqueous phase in neutralized by transfer of at least a portion of thebasic moiety to the aqueous phase. The extractant phase is made up ofone or more of an cationic extractant compound, most preferably one ormore of a oil soluble quaternary ammonium compound, comprising R₄N⁺, butalso could be an alkylated guanidinium compound, R₅CN₃H⁺, or analkylated quaternary phosphonium compound, R₄P⁺, or a blend of thesecationic compounds; wherein the associated anion is a pH basiccomponent. Typically the equilibrium pH of the unstable emulsion iscontrollable from about 2 to about 13; the equilibrium pH of theunstable emulsion is typically about 3 to about 9, or from 4 to about 11depending on desired ionic components for removal. This pH control isachieved by varying the E/A ratio, with pH increasing with this ratio.Also, the extractant alkyl groups can also be aryl, and/or alkylaryl.

If desired, more preferably, one or more of an optional hydrocarbondiluent and/or a modifier, most preferably nonflammable, for promotingthe aqueous and extractant phases to remain fluid, disengage and dewaterin a following settling step, and the extractant phase homogenate, maybe added as part of the extractant phase; disengaging mixture of atreated aqueous phase and a component-loaded extractant phase from theunstable emulsion, wherein a new stable emulsion of floc phase forms inthe loaded extractant of a new composition of matter phase, and whereinthe extractant phase and aqueous phase are at least partially separatedby allowing gravity separation, optionally sped up by centrifugationand/or use with a hydrocyclone; and where the floc remains in theE-phase but may tend to settle to the interface with the aqueous phase,and separating the loaded extractant phase and floc from the aqueousphase by using a skimming weir design to ensure that the floc is removedfrom the aqueous phase along with the E-phase and before it thickensinto a slow to non-flowing solid precipitate.

The R groups in the quaternary ammonium compound, the guanadiniumcompound, or the quaternary phosphonium compound can be differenthydrocarbons that have a carbon atom content of at least one each andwith a total carbon number for the R₄N⁺ compound of and/or R₄P⁺ compound37 carbon atoms with a minimum total carbon number of nine, butpreferably 25. The basic requirement is that the “R” alkyl, alkylaryland/aryl groups provide sufficient hydrophobic character to theextractant phase so that it remains a separate phase from the aqueousphase whether loaded with floc or basic moiety. One particularly usefulquaternary ammonium compound, N-methyl N,N,N-tri(n-octyl)ammonium ion,used in the examples herein, has alkyl groups with a straight chainlength of one carbon for one R group and a chain length of eight for theother three R groups. The R groups may be the same or different and bearomatic, aliphatic, or mixed aromatic/aliphatic. The R groups cancontain other groups or other atoms of Si, F, Cl, Br, O, N, S, or P solong as they do not make the ionic compound too oil insoluble, too watersoluble, or neutralize or change its electronic charge. Therefore,halogens such as chlorine and the like, halogenoids, ethers, esters,imines, ketones, phosphate esters, nitriles, and the like arepermissible.

The extraction phase to aqueous phase ratio (E/A) is typically of about1:20 to about 20:1 in the extraction circuit, preferably 2:1 to 1:10 andmore preferably 1:3 to 1:7, and most preferably 1:4 to 1:6. Theextractant concentration in the E-phase needs to be at least 0.1% andcan be neat (100%), more preferably 2-30% and most preferably 3-15%.

The pH of acid mine water is typically about 4-7 while in the mine andnot exposed to air. On seeping or flowing from the mine, or otherwiseexposed to air, the pH typically drops to about 2 to 4 depending on thedissolved metals in the water, especially when ferrous ion, ferric ion,and aluminum ion are present in the AMD water, normally the case.

The invention provides efficient means to purify water, especially theremoval of anions and cations present in the feed water contributing tothe waters' total dissolved solids level, collectively referred to astotal dissolved solids solutes, values and/or contaminants. Thoughapplicable for purifying water at all scales, most preferably theinvention is useful in purifying very large volumes of water of tens toten thousand gallons per minute flow rate in a continuous-flow fashionusing a unique combination of fast kinetics process chemistry that makesuse of interacting thermodynamic driving forces of ion pairing,acid-base chemistry, gas evolution, and physical separation technologiesbased on flotation (F) and liquid-liquid extraction (LLX) technologiesapplied in liquid-liquid skimming configuration, and therefore, isreferred to as flotation liquid-liquid extraction process technology.Flotation liquid-liquid extraction technology of the invention providesa major array of benefits including:

Fast removal of water contaminants in seconds to minutes in total or pertreatment stage, rather than hours or days and weeks required byconventional approaches.

Fewer water contacting stages (where it is noted that, especially forhigh flow rate feed waters, the larger the number of stages required topurify the water, the more costly such treatment becomes since the samelarge volume of water needs to be retreated at each stage),

Removal of metal ions normally not extractable or not well extracted byconventional liquid-liquid extraction (LLX) methods, and/or which foulconventional LLX processes due to in-process solids precipitation,especially Fe, Al, and Ni.

Extraction of a broader spectrum of ions, anionic and cationic, thanconventional methods.

Extraction of anions including sulfate, selenate, arsenate, molybdate,vanadate, gold halo/pseudohalo complexes, chromate(VI), permanganate,and the like, and cations including Fe(III), Fe(II), Cr(III), Cu(II/I),Ni, Co(II), Co(III), Zn, Al, Mn, Mg, Ca, La, Ga, Cd, Hg, CH₃Hg⁺, Ag(I),Au(I&III) and the like, to low residual concentrations rendering thewater suitable for potable water production, agricultural use,industrial process water, and the like.

Concentration rather than dilution of contaminants with preferredseparation of the extracted ions into product concentrates, withoutbeing bound by low value waste minerals, normally forming concentratessuch as metal ion sulfate solutions, solid salts, and/or solids ofoxides, hydroxides, bicarbonates, and/or carbonates, allowing them to beused to make useful and saleable conventional products such as metalproduction (iron, zinc, cobalt, nickel, copper, aluminum, and the like),sulfate-based fertilizers, especially potassium sulfate, ammoniumsulfate, soil acidulants, and/or sodium sulfate, a unique composition ofmatter consisting of a combination of ferric, ferrous, or aluminumsulfates, and the like, with many applications.

Low-cost means to purify large volumes of water continuously, in asimple manner, with a high yield of purified water, and with arelatively small equipment size when compared to conventionaltechnologies.

Concentration of cations and anions in one or more liquid-liquidextraction stages without formation of solid product waste sludges suchas precipitants, gypsum, lime sludges, or the like.

Description of Acid Mine Waste Drainage Water Purification Processwithout Waste Generation and with Metals and Sulfate Product (Values)Recovery.

The invention includes a second new composition of matter materialconsisting of a unique formulation containing at least one liquidextractant component combined with a metal ion colloid consisting of onemetal ion in neutral combination with anions of oxide, hydroxide,carbonate, sulfate and/or bicarbonate ions that, when combined, form acolloid, floc, or suspended particulate, that is useful in thesimultaneous and rapid separation of metal ions and anions from verylarge volumes of water at low cost and without waste generation.

The water insoluble floc produced in the extractors of the F-LLX processof the invention in the pH range of 5 to 8 is known to consist of thefollowing:

-   -   metal ions derived from the AMD water contaminants, especially        Fe and Al,    -   cationic extractant derived from the extractant phase,    -   hydroxide ion derived from the reaction of carbonate ion with        water during the extraction operation, and    -   sulfate ion removed from the AMD water by the extractant phase.

This brown floc is less dense than water despite it containing inorganiccomponents. It can be isolated by filtering or preferably bycentrifugation after separating it from the two phase contents of thesettlers and decanters. The E1D decanter is the primary apparatus andoperation to prepare this new floc material since the floc generatedthere is in largest quantity and contains the least water and diluent.The floc solid is useful in purifying water by enabling removal of TDSfrom contaminated waters and as a flexible means for preparing metal ionsalts and alkali salts of sulfate ion but without the introduction ofother salts that is common to these synthesis methods since, with thefloc, the counter ion is the extractant. Although carbonate ion loadedextractant phase is used to raise the pH of the AMD water, as anexample, carbonate ion itself is not a component of the oil soluble flocpreviously described, rather it reacts with water by hydrolysis togenerate hydroxide ions, that forms metal hydroxide colloids (seebelow), and it forms certain insoluble metal ion carbonate particulatesthat are not part of the floc (see below). Hence, since the metal ionsand sulfate ions are removed from the water phase when the floc isformed at pH 5 to 9, the floc empirical formula is estimated to be,

(R₄N⁺)_(x)(M(II))_(n)(OH⁻)_(z)(SO₄ ²⁻)_(w) (for the case of divalentmetal ions, M(II))  (Eq. A)

or

(R₄N⁺)_(x)(M(III))_(y)(OH⁻)_(z)(SO₄ ²⁻)_(w) (for the case of trivalentmetal ions, M(III)  (Eq. B)

-   -   (Where M(III) can be Fe(III) or Al, and M(II) can be Fe(II), Ni,        Co(II), Cu, Zn, Pb, Cd, or Mn)        Now, using charge balance, the following relationships hold,

x+n=z+w, for the M case  (c)

x+y=z+w, for the N case  (d)

AMD water analytical results provided in Table 2A were used with theabove observed descriptive and chemical analyses to determine the valuesof x, n, y, z and w. These results are given, using the Fe(II) andFe(III) cases, as follows:

(R₄N⁺)_(x)(Fe²⁺)_(y)(OH⁻)_(z)(SO₄ ²⁻)_(w)  Floc Chemical Composition:

-   -   For which, x=2-3, y=1-0, z=2 and w=1-2

And

(R₄N⁺)_(x)(Fe³⁺)_(y)(OH⁻)_(z)(SO₄ ²⁻)_(w)  Floc Chemical Composition:

-   -   x=2-3, y=1-2, z=2 and w=1-2        Results in Table 2B generated the values of x, y, z, and w for        the following extraction phase as follow:

(R₄N⁺)_(x)(Fe²⁺)_(n)(OH⁻)_(z)(SO₄ ²⁻)_(w)

-   -   x=2-4, n=1-0, z=2 and w=1-3

(R₄N⁺)_(x)(Fe³⁺)_(y)(OH⁻)_(z)(SO₄ ²⁻)_(w)

-   -   x=2-4, y=1-2, z=2 and w=1-3        For highest yields of these materials, it is important that the        quaternary ammonium ion is used in excess or at least in charge        balance as calculated using the above equations and for the AMD        water composition of cations.

This material is insoluble in water and is oil phase dispersable, it cancontain solvent molecules of water and/or alcohols and/or esters ofcarbon number of at least eight, preferably 8 to 24, and most preferably10 to 16, so long as the quantities of solvent do not collapse thecolloidal material or render in water soluble. The invention is a newcomposition of matter useful for the low-cost purification of acid mineor acid rock drainage or ground waters, natural gas well brines,extractive metallurgical aqueous processing waters, surface finishingprocess waters, agricultural processing and waste waters, and the like.It is especially useful to reduce high metal ion and sulfate ionconcentrations in water, for example as is found in mining and mineralprocessing effluent waters, in brackish and brine waters, in industrialprocess effluent waters, livestock farm waters, and the like, down fromvery high levels, for example total dissolved solids values of 300-5000mg/L, down to levels that allow water re-use, for example to <500 mg/Ltotal dissolved solids for surface water discharge, or <250 ppm toenable re-use as drinking/potable water, and/or to deionized water withtotal dissolved solids of <125 ppm in which most or all of the containedtoxic metal ions are removed to <1 or normally <0.3 ppm, and sulfate ionto <30 ppm, phosphate ion to <10 pppm, and nitrate ion to <10 ppm. It isbelieved that this efficient deionization is accomplished by a uniquephysio-chemical process based formation of a new hydrophobic colloidalmaterial consisting of a oil soluble quarternary ammonium compound orblend of such compounds (and/or a oil soluble phosphonium compound orblend of such compounds) combined with a blend with of at least one, andpreferably more than one type of metal ions with anions selected fromthe list of oxide, carbonate bicarbonate, hydroxide ions, sulfate ion,phosphate ion, and/or nitrate ions, and where the colloidal material canbe in suspended precipitate form, but most preferrably it is in flocform, and most preferrably in colloidal form. The colloid, floc, orsuspended precipitate can optionally also contain up to 70% entrainedwater, preferrably less than 25% water, and more preferrably <10% water,and most preferable <7% water. In addition, the colloid, floc, orsuspended particle and/or precipitate can also contain up to 80% oilsoluble alcohol, ester, alkyl phenol, and/or other modifier, and mostpreferably also contains a diluent. All of the organic extractant phasematerials are preferred to be nonflammable compounds.

As is known in the art, sulfate ion is very difficult to remove fromwater regardless of purification technology selected for the importantniche of medium to high flow rates when present at medium to high levelsof SO₄ ⁼ in the range of about 500 to 10,000 mg/L. The need is to reducethe sulfate ion concentration far enough to enable surface dischargerelease of the water, into the range of 0 to <600 ppm needed for surfacedischarge release, and preferably to enable its use for potable water(<250 ppm), and most desirable to <100 ppm needed to prevent drinkingwater “sulfur” odors. This difficulty in removing sulfate ion is wellknown in the art of water purification and is due to sulfate ion's highcharge density and high water solubility of all its salts except bariumion, and barium ion is very toxic, very expensive, very difficult torecycle, and is in severely limited supply relative to the amount ofwater needed to be purified with respect to sulfate ion. Thecommercially available technologies all have the a major problem in thata substantial amount, normally 25 to 50%, of the water to be processedends up as a waste stream carrying the contaminants from the process toregenerate the sorbent or avoiding scaling of the process hardware.Hence the technologies of the known art produce a large waste streamthat is still of a very large volume, still dilute, still toxic, andstill represents a severe and large disposal problem.

In the present invention, a unique “forced” ion pairing processchemistry is used to quickly remove highly soluble and/or highly chargedanions, especially SO₄ ⁼ to <100 mg/L residuals, and in high yield fromthese intermediately concentrated (about 600 to 10,000 mg/L totaldissolved solids feedwaters. The invention maximizes the use ofoff-the-shelf hardware to help speed up the availability and widespreaduse of this much needed technology.

The process chemistry of the invention involves contacting (mixing)contaminated aqueous solutions with a specially formulated,nonflammable, water immiscible, oil soluble, extractant phase. Thesulfate is extracted into this second liquid phase, that is also used ata volume (or flow rate) that is substantially less than that of thecontaminated feedwater so that the sulfate and other ions co-extractedare pre-concentrated by the factor of the volume (flow rate) ratio. Inthis manner the water impurities are not only removed from the water butare simultaneously concentrated multiple times so that a concentratedmetal sulfate product of saleable value is produced rather thanproduction of wastes. To boost process efficiency and to lower processcosts, an aspect of the invention is that, at the same time the sulfateion is extracted, the toxic metal ions contained in the feed water arealso removed and concentrated using the same operation. The capabilityof the extractant phase to extract both sulfate ions and metal ionssimultaneously into a water immiscible phase is key to the invention andoccurs due to the process chemistry and process operating conditionsselected for the invention.

This colloid/floc process chemistry mentioned above is described furtherbelow.

The extractant chemistry discovered for the role of SO₄ ⁼ removal is aformulation containing an oil soluble quaternary amine {R₄N⁺}. The useof quaternary ammonium LLX extractants for anion extraction is known inthe prior art. However, this art has consistently indicated that sulfateion is only poorly extracted by these reagents and far prefersextracting mono-anions and anionic metal ion complexes instead.Therefore it was necessary to modify the process chemistry such topromote sulfate ion partitioning into the water immiscible extractantphase. This “forced” approach to the process chemistry will now bedescribed.

Choice of Anion and Base for pH Control

Since acid mine drainage waters and the like are already very highlyexcessive in total dissolved solids level, simple ion exchange removalof sulfate ion, for example by using the chloride ion form of thequaternary ammonium extractant, was out of the question since any anionexchanged for the sulfate ion would then render the water still a wastetoo toxic for discharge to surface waters, much less be of any value foruse by industry or residential. Therefore anions that can be eliminatedby forming H₂O and/or a gas, specifically CO₂ gas, were used to developthe process chemistry, that is the hydroxide ion (OH⁻), carbonate ion(CO₃ ⁼), carbonic acid (H₂CO₃), and/or bicarbonate ion (HCO₃ ⁻) form ofthe quaternary ammonium extractant}. It is also important that thequaternary amine does not require protonation to possess a positivecharge and is positively charged over the entire pH range. In thismanner the strong base options of carbonate ion and/or hydroxide ion canbe used. The process chemistry identification and selection is furtherdeveloped below to explain the fundamental separation process chemistrythe proposed mechanism of action.

Preferred Extractant Structure and Means of Process pH Control

A quaternary amine liquid-liquid extractant particularly suitable topractice this invention is N-methyl tri(octanyl) ammonium ion availableas Aliquat® 134 or Aliquat® 336 (Cognis, Inc.), or oil soluble, lowwater soluble quaternary ammonium compound, preferably a liquidcompound, and most preferably a branched or extensive tripodal structureto discourage gelation and solidification throughout the process load,strip and storage cycle.

Description of Colloid Formation and its Impact on Separation ProcessChemistry Basis

One embodiment describing the invention is the general process flowdiagram FIG. 1A using acid mine drainage water as the “Metal SulfateLaden Feed Water” 1110 example for the description. The first stepinvolves treatment of the acid mine drainage water stream with aparticular Aliquat 134-based extractant phase formulation 1010 in aself-controlled pH buffer system of carbonate/bicarbonate buffer,included in the extractant phase formulation with the Aliquat 134, andcontrolled in the preferred range of pH 5.5 and 9.7 for at least onestage, preferably two stages, and where up to five stages is effectiveand 4 stages with the preferred pH range for extraction stage 2030ranges from 7.4 to 8.2 and is preferably 7.6 to 8.0 and is mostpreferably about 7.8. The stages can be configured entirelycounter-current, or cross-current where one or more sulfate-strippedextractant phases (see below) are blended with the acid mine drainagewater as it flows serially through two or more stages, or co-currentflow, cross-flow, or a combination of these flow configurations. Mostpreferred is to operate the extractors with the four stages separatedinto two sets of at least two mixer settlers each, where each set isarranged counter-current with its partner, and each set internally ispiped up to be counter-current, but the two sets are configuredcross-current flow with respect to each other where the extractant phaseis the crossing phase and the aqueous phase is piped to flow seriallyfrom stage to stage across the first set and then across the second set,and then, if present, across the third set, and so on. However, othercombinations are easily arranged and tested since piping changes arereadily made as is already well appreciated in the current art ofliquid-liquid extraction technology using staged devices such asmixer-settlers, centrifuges, columns, in-line static mixers, and thelike and any combination of these. Most preferred is the, in this caseis a combination of counter-current and cross-flow configurations,described above, while operating the strippers always withcounter-current flow and internally aqueous recycled. Also, mostpreferred to the invention, is that the extractant phase approaches inthe settlers are specially designed for floc handling by promotingsimultaneous floc, colloid, floating precipitate, extractant liquidphase, and/or gelling mass densification, while still maintaining itscontinuous, or at least intermittent, flow and by promoting the flow, atleast intermittent and preferably continuous, of this floc over theextractant phase exit weir, by any combination of mechanical movementand/or smooth, curved surface as, for example, is found in ore flotationcell designs, and lastly by accommodating the flow requirements of thisextractant phase “floc”, preferably thickened, to the next stage, wherethe flow can be intermittent or continuous, whether it be anotherextractor stage, a decanter for further dewatering, or a stripper. Thereceiving hardware for the floc from the flotation overflow weir can beof hopper design to send the floc to the next stage, a pipe ofpreferably relatively large ID, or a trough, a hydrocyclone, or anycombination of these.

The most preferred case is that the following stripper uses an aqueousacid to dissolve the components of the extractant-based floc, whether itbe a colloid, precipitate, slurry, liquid, emulsion, gelling mass, orany combination of these, thereby returning the system back into aliquid-only condition. Liquid-only systems are easiest to handle andprocess under continuous flow.

Suggested Mechanism of Action

While not wishing to be bound by theory, but acknowledging that theoriesare useful in understanding physical/chemical processes, theinterrelationships of control parameters with process outputs,capabilities of the process, and the like, it is presently believed thatcontacting contaminated water, preferably containing dissolved mineralsof cations and anions, with the invention extractant phase method ofoperation, and hardware generates a hydrophobic and variable combinationof oxo and hydroxo, optionally in combination with carbonato,bicarbonato, sulfato and/or other metal on ligand components and anionspresent, with metal ions containing homogenous, but usuallyheterogenous, material with the physical consistency, depending on age,temperature and concentration, with the properties of one or more of thefollowing: metal ion complexes, clusters, colloids, flocculatedparticulates, floating precipitates, liquids and/or gels, and that canhave various degrees of hydration from low to high, and will becollectively referred to as a “floc” in this description.Microscopically, it is believed that this floc is essentially neutralbut preferably slightly anionic so that it interacts with a positivelycharged hydrophobic “extractant” dissolved in a water immiscible liquid.At the molecular level, it is believed that this floc forms quickly,within a few seconds since metal ion to ligand bonds need not be brokenor formed. It is believed that this capability is an important featurethat makes the invention immensely useful for cheaply purifying largewater flows of a broad spectrum of contaminants. The floc consists ofoxo, hydroxo, with or without other anions, providing colloid formingcomplexes of the metal ions present, for acid mine drainage watersnormally one or more of Fe(II), Fe(III), Al, Mn(II), Mn(IV), Ni, Co(II),Co(III), Zn, Mg, Cu, and Ca. As provided by the inventive formulation,simultaneously sulfated quaternary amine, (R₄N)₂SO₄, also forms andremoves sulfate ion from the water in high yield, and, in another aspectof the inventive formulation, simultaneously neutralizing the acidpresent in the acid mine drainage water by the conversion of H⁺ ionspresent and provided by acidic metal ions into H₂O and H₂CO₃ moleculesby neutralization with CO₃ ⁼ and/or HCO₃ ⁻ ions to make H₂CO₃, and, inanother aspect of the inventive formulation, simultaneously reducing thetotal dissolved solids level of the water through the dissociation ofH₂CO₃ into CO₂ gas and H₂O, thereby producing a purified aqueous phase(water). Selenate and silica/silicate are also recovered. The metalsrecoverable include those capable of metal hydroxide precipitations,metal carbonate solid formation, metal sulfate solid formation, or thosemetals that co-precipitate by inclusion in to ferric hydroxide and/oraluminum hydroxide amorphous precipitates, and/or those anions that formhighly insoluble precipitates with ferric ion or aluminum ion, such asarsenate, phosphates, humic acids, other natural organic matter, and thelike. This includes all the metal ions on the Periodic Table exceptalkali metal ions, which are reduced only a low to small amount. Alkalimetal ions are typically a minor component of ground waters, surfacewaters, farm waters, and most industrial process waters. The low removalof alkali metal ions is not a detriment since in fact there is muchallowance by water users that these ions can remain in the productpurified water, especially the nutrient K⁺ and where softened water isused (where Na⁺ residuals exist in the treated water). For this reasonalso, the technology of the invention is not useful for the desalinationof sea water where the main objective is the removal of highconcentrations of the alkali salt NaCl is required.

Metal ions removed by the process typically constitute those metals thatform oxide, hydroxide, or carbonate colloids, and those thatco-precipitate with such materials (FIG. 2). Also attached is Table 3summarizing the list of effective extractant structures. These can beused in water immiscible solution form. They can be used neat if theircorresponding bicarbonate/carbonate/bisulfate or sulfate ion forms aresufficiently fluid at short mixing and settling times (see below) so toenable flow in batch or continuous liquid-liquid separation process orflotation process or oil skimming circuits. Note that the prior artteaches that all such metal hydroxide and insoluble salt compositionsystems will form solids which precipitate within conventional LLXhardware requiring process hardware shutdown, solids removal, repair ofdamaged equipment and other maintenance. The current invention avoidsthis serious problem. If enhanced fluidity is needed as is normally thecase, then one or more of the quaternary extraction compounds areblended with a predominantly hydrocarbon diluent of eight or more carbonatoms. Suitable diluents are included in Table 3. Also included in Table3 are candidate “modifiers” that can be added to the extractionformulation that can aid in displacing entrained water in the extractantphase carrying the sulfate ion and dispersible metal ion complexesand/or colloids of oxo, hydroxo, carbonato, and/or sulfato extractantphase flocs. Modifiers can also help the floc to solubilize in thehydrocarbon diluent.

Incorporation of Floc Filtration Option

In another embodiment of the invention, the metal ion/sulfate ion flocmaterial, loaded onto the extractant phase during the mixing step (FIG.1C) is filtered after the extraction operation and/or in betweenextraction stages, to remove the novel composition (described above)consisting of the formulas:

(R₄N⁺)_(x)(M(II))_(n)(OH⁻)_(z)(SO₄ ²⁻)_(w) (for the case of divalentmetal ions, M(II))  (a)

and

(R₄N⁺)_(x)(M(III))_(y)(OH⁻)_(z)(SO₄ ²⁻)_(w) (for the case of trivalentmetal ions, M(III))  (b)

This filtration illustrates that the new material described can beproduced as an isolated product. This filter step is optional and wouldbe performed before the metal ion and sulfate ion stripping steps (seebelow). The filtrate, now essentially 100% liquid colloidal floc andliquid, is then processed using liquid-liquid contactors in thestripping of soluble metal ions and of sulfate as described below.Sulfate Ion Product Production with Concomitant Regeneration ofCarbonate Form of the Extractant Phase

After liquid-liquid stripping and/or filtering the metal ion productsfrom the extractant phase, as described above and in the attachments andexamples below, the essentially metal ion-free extractant phase stillcontains the sulfate ion (and any selenate ion) recovered from the acidmine drainage feed water during the extraction operation. The chemicalform of the sulfate ion on the extractant phase is believed to be asionic compounds (R₄N⁺)₂SO₄ ⁼ and/or R₄N⁺HSO₄ ⁻ depending on the pH ofthe extraction or metal ion stripping stage operation stage from whencein came, normally the former from the extraction circuit and the latterfrom the acidic metal ion strip circuit, is contacted in counter-currentLLX configuration with 0.1-55.% Na₂CO₃ or 0.1-55% K₂CO₃ solution, orsolutions of bicarbonate ion, or blends of these, with or without addedNaOH or KOH up to saturated conditions and could be optionally warmed toachieve even higher extractant phase loadings of these anions. Totalconcentrations of carbonates than solubility limits can be used providedthe resultant slurries are kept in flow motion, at least intermittently,by mechanical means of stirrers and pumps and use of troughs and largepiping instead of small ID piping typical of conventional LLXtechnology. With use of suitable solids handling equipment, i.e. hopperwith auger addition trough, solid carbonate and hydroxide solids can beadded directly to the sulfate strip mixers as powders, pellets,granules, and the like. With this information, it would be obvious toothers skilled in the art of conveying slurries to prepare relatedmechanical designs for the physical handling of the product “floc” ofthe invention as described above.

At the preferred counter-current arrangement of the sulfate stripperoperation the “first” sulfate strip stage, “S1-SO4”, generates anaqueous raffinate from the sulfate stripper stage that is the mostconcentrated in sulfate ion and represents the “sulfate productconcentrate”. Depending on relative flow rates of the extractant phaseand the carbonate/hydroxide strip feed solution, the sulfate ion productconcentration can be adjusted over a very wide range of approximately2,000-650,000 mgSO₄ ⁼/L. Preferably the product sulfate ion concentrateis 150,000-250,000 mg/L (or about 15-25%) for the case of stripping withNa₂CO₃, and 5,000-150,000 mgSO₄ ⁼/L for the case of stripping withK₂CO₃.

Although in the prior art the anion-anion liquid-liquid exchangeinvolving sulfate ion is normally achieved with poor efficiency andincompletion in the case of extracting sulfate ion from water, thenature of the invented process disclosed herein provides sharp, fast andhigh yield sulfate ion recovery and stripping in stages Sy-SO4 (wherey=1 to 3 stages, preferably 1 to 4 stages, and most preferably 1 to 5stages, while 1 to 6 stages is also effective), by providing thecombination of strong thermodynamic driving forces of acid/baseneutralization {favorable (negative) enthalpy change} and innocuous gas(CO₂) formation {favorable (positive) entropy change}, and low E/Aratio, dianion exchange, and release (positive mass action effects),summarized by the following equations for sulfate ion (equivalentreactions can be written for nitrate, chloride, methane or otherssulfonate(s), phosphate, acetate, and other anions that represent theconjugate base of the acid used in the metal ion strip circuit usedseparately or in any combination), and where Extractant Phase=E-phase,and the undesignated phase is Aqueous Phase,

2{R₄N⁺HSO₄ ⁻}_(E-phase)+M_(x)CO₃→H₂O+CO₂(g)+{(R₄N⁺)₂SO₄ ⁼}E-PhaseM_(x)SO₄  (1)

For example, for x=2 and SO₄ ⁼ strip stage pH of ˜2<pH<˜7,

2{R₄N⁺HSO₄ ⁼}_(E-phase)+K₂CO₃→H₂O+CO₂(g)+{(R₄N⁺)₂SO₄⁼}_(E-Phase)+K₂SO₄  (1′)

and/or,

2{R₄N⁺HSO₄ ⁻}_(E-phase)+2M_(x)CO₃→2M(HCO₃)_(x)+{(R₄N⁺)₂SO₄⁼}_(E-phase)+M_(x)SO₄  (2)

For example for x=2 and SO₄ ⁼ strip stage pH of ˜8.5<pH<˜10,

2{R₄N⁺HSO₄ ⁻}_(E-phase)+2K₂CO₃→2KHCO₃+{(R₄N⁺)₂SO₄⁼}_(E-phase)+K₂SO₄  (2′)

and/or

{(R₄N)₂SO₄}_(E-phase)+M_(x)CO₃→{(R₄N)₂CO₃}_(E-phase)+M_(x)SO₄  (3)

As an example, for x=2 and SO₄ ⁼ strip stage pH of >10,

{(R₄N)₂SO₄}_(E-phase)+K₂CO₃→{(R₄N)₂CO₃}_(E-phase)+K₂SO₄  (3′)

and/or

{(R₄N)₂SO₄}_(E-phase)+2M(HCO₃)_(x)→2{R₄NHCO₃}_(E-phase)+M_(x)SO₄  (4)

Where x is 2, and where the {R₄N⁺HSO₄ ⁻}_(E-Phase) and/or{(R₄N)₂SO₄}_(E-phase) species flow into the sulfate ion strip circuitfrom the metal ion strip circuit(s) and the anion is that of the stripacid used in the metal sulfate stripping, and for acid mine drainagewater feed may also contain sulfate/bisulfate ion. The relativecontribution of each of the above four reactions at each stage ofstripping depends upon the pH of operation of the last stage of themetal ion stripper circuit (or of optional wash stage(s)), that is,whether the anion on the extractant phase is sulfate ion, bisulfate ion,or a blend of these, and on the pH of each stage of the sulfate stripcircuit. The use of pH to control the strip circuits, including morespecific effective pH values, is further detailed below.

For the above, M is most preferably Na⁺ or K⁺ (x=2), or blends thereof.These ions can be used alone or combined with any of the following: NH₄⁺ (x=2) provided the pH is <˜9) and especially as concentrated solutionsof ammonium bicarbonate (NH₄HCO₃) to produce concentrated productsolutions of ammonium sulfate, ammonium nitrate, ammonium phosphate, andthe like and/or blends thereof. Given the above information, it isobvious to those skilled in the art of liquid-liquid extraction that thebasic anion can also be added back onto the extractant phase after thestripping of the sulfate (or other anion) by a neutral anion, e.g.sodium or potassium nitrate and/or phosphate solution, followed by aseparate sequence of contactors to replace these ions with carbonateion, or the like. This mode of operation however requires additionalcontactors and chemical raw materials, and is therefore less preferred.

The generation of CO₂ gas in this manner is a unique feature of theinvention and occurs most when the last stage of metal sulfate stripping(S3M in the most preferred case), or wash stage (also in the mostpreferred case) operations are acidic (pH<3, and most preferably pH<1),as it is in the case where the strip acid conjugate base is a weak acid,as it is in the case of sulfuric acid (bisulfate conjugate base),orthophosphoric acid (orthophosphate, monobasic being the conjugatebase). It does not occur significantly when the conjugate base of thestrip acid is not a weak acid, for example when nitric acid orhydrochloric, other hydrohalic, MSA (methane sulfonic acid), and thelike.

The CO₂ release occurs preferentially in stage the first stage of anionstripping. For the sulfate case this is stage S1-SO4 (first stage ofcarbonate/bicarbonate stripping of bisulfate ion) and is due to Reaction1 where the control pH at the S1-SO4 mixer is most preferably 4.5±1,preferably 4.5±2, but control in the pH window of 2 to 7 is stilleffective. The CO₂ so released is humid, but otherwise pure if theS1-SO4 mixer is reasonably sealed against air intrusion. This CO₂ gas isideal for adjusting the pH of the finished water back down to <9 whenthe final extraction stage is operated at pH 9-11 to remove Mg and Ca tosoften the product water and to further reduce total dissolved solids(see below). Being pure, the CO₂ gas can be captured, pressurized andpackaged as liquid or dry ice by current commercial and well establishedtechnology to produce a useable CO₂ product. Otherwise it can beharmlessly vented. This CO₂ gas does produce a rapidly breaking foam inthe S1-SO4 mixer that needs to be handled such to prevent tank overflowspillage during the operation. If the E/A ratio in the stripper isinsufficient in phase, that is the pH in S1-SO4 stage exceeds ˜7, thensome foaming may aqueous also occur in stage S2-SO4. This foaming isfound to be a minor issue as the foam breaks rapidly and only a smallfoam head is produced. For this reason a continuous stirred mixer tankis most preferred with higher walls on the mixer and associated settlercompartment of the first and second carbonate strip stages compared toconventional liquid-liquid technology (see below). Most preferred isthat S1-SO4⁼ also contains a cover that enables the capture of the CO₂gas, without interfering with the mixer, to use to sparge through thepurified product water (see below), if needed, to optionally adjust itspH downwards, preferably to pH<10, and most preferably to pH 9.

When the aqueous carbonate solution being flowed to the S1-SO4 sulfatestripper stage is derived from a sulfate ion stripper stage just upstream of the S1-SO4 stage in a counter-current operation, i.e. S2-SO4,the most preferred case, then some or most of the carbonate may alreadybe in the bicarbonate ion form (see above chemical equations), i.e.M_(y)HCO₃, where y=1 for M=Li⁺, Na⁺, K⁺ and/or NH₄ ⁺, and y=0.5 forM=Mg²⁺, Ca²⁺ and/or Zn²⁺. If M=NH₄ ⁺, as is well known in the art ofhandling ammonia aqueous solutions, that the form of a very solublereagent addition is NH₄HCO₃, prepared from a solid and/or concentratedsolution, or directly from NH₃(g) gas or aqueous ammoniacal solution,NH₃(aq).

The extractant phase from S1-SO4 can still contain about one half of thesulfate ion from the extraction circuit depending upon the overall E/Aratio of the stripper circuit and the concentration of active quaternaryammonium concentration used in the extractant, and the E/A flow ratioused in the extraction circuit. This phase is then contacted, preferablyin counter-current liquid-liquid extraction mode using suitable hardware(e.g. mixer-settlers, static in-line mixers, columns, or continuousliquid-liquid centrifuges, or hydrocyclones), at least one additionalcontact time, preferably two to three more, and most preferably four tosix more times; that is to say using additional counter-currentconfigured stages, for example S2-SO4, S3-SO4, and S4-SO4 for a total offour anion (sulfate) strip stages linked counter-current. Hence theextractant phase exiting the “last stage” of the sulfate ion stripcircuit (normally S4-SO4, but could be S3-SO4, S5-SO4 or S6-SO4), is thecarbonate ion-loaded extractant phase (i.e. the form containing{(R₄N⁺)₂CO₃ ⁼}_(E-phase)) that is recycled to the extraction circuit(e.g. see attached detailed process flow diagrams).

If hydroxide ion, OH⁻, is to be also added to the CO₃ ⁼-loaded E-phaseto reduce the amount of residual total dissolved solids by forming lessHCO₃ ⁻ in the purified water and/or by removing hardness metal ions,then OH⁻ is added to the SO₄ ⁼ strip stage mixer (normally to stagesS3-SO4, S4-SO4, S5-SO4, and/or S6-SO4 depending on the desired (SO₄ ⁼)residual desired in the treated water and/or on the E/A used in theanion stripper operation).

The raffinate phase flowing from the S1-SO4 stage, either continuouslyor intermittently, is the useable sulfate ion product, for exampleconcentrated solutions and/or easily crystallized solids of (NH₄)₂SO₄,K₂SO₄, and/or Na₂SO₄ (as a hydrate or anhydrous form). These productsare in addition to the metal salt solution products or solids, forexample metal sulfates produced from the S1-M metal ion strip stagecircuit referred to above, or from both S1-M and S1-N stages from a dualmetal ion strip circuit (see below). These products are items ofcommerce used in numerous industries, with the metal salt concentrateblend being of a new composition (see above).

As illustrated by the discussion above, and the Detailed Description ofthe Invention and Examples sections, the exact pH and E/A flow ratios ofeach flotation liquid-liquid extraction contactor, the added reagentsand reagent concentrations, and the like, used determine the purity ofthe water produced and the type and amount of co-products produced. Thisflexibility of products help insure that the product mix and puritiesproduced can be best tuned to meet regional municipal, industrial,agricultural, and residential market demands for the products producedin practicing the invention.

Operation of the Sulfate Recovery Circuit and Regeneration of CarbonateForm of the Extractant Phase for Continuous Recycle

In a most preferred embodiment, if the feed of acidic extractant phaseto S1-SO4 is in excess of the molar amount that can be handled by theabove chemical reactions 1 to 4, then all of the aqueous carbonate isconsumed, as desired, in S1-SO4, as given in the above chemicalequation, thus insuring that no carbonate exits the process with thesulfate product and thereby insuring that all the carbonate added to thesystem is transferred to making sulfate concentrate product (or otheranion product as listed above). In the case of some excess acid beingpresent in S1-SO4 (indicated by a S1-SO4 mixer or raffinate pH of <4),(the acidity being supplied from the last metal ion stripper stage,normally S3M, but optionally could be S2M or S4M), and present as{R₄N⁺HSO₄}_(E-phase), or as {R₄N⁺H₂PO₄}_(E-phase) and/or {R₄N⁺HPO₄⁼}_(E-phase) if phosphoric acid is the strip acid used, still exists onthe extractant phase and this is easily consumed, and the associatedsulfate removed, in the S2-SO4 stage, but CO₂(g) does not form, asfollows (shown for the sulfate ion system):

{R₄N⁺HSO₄}_(E-phase)+M_(x)CO₃→M_(y)(HCO₃)_(z)+{(R₄N⁺)₂SO₄}_(E-phase)  (5).

and in parallel,

{R₄N⁺HSO₄ ⁻}_(E-phase)+M_(x)CO₃→M_(x)SO₄+2{R₄N⁺HCO₃ ⁻}_(E-phase)  (6)

Where the M, x, y and z definitions are as before. In this case, theextractant phase carries the {(R₄N⁺)₂SO₄ ⁼ }_(E-phase) and {R₄N⁺HCO₃⁻}_(E-phase) species to the S3-SO4 and/or S4-SO4 strip stages where theSO₄ ⁼ and HCO₃ ⁻ extractant phase species are displaced (ion exchangedout) by the overwhelming high concentration of divalent (pH 10+) CO₃ ⁻ion. The carbonate ion concentration used is limited only by thesolubility of the carbonate salt and, more preferably, also by theproduct of the salt solution in stage S1-SO4 (again, using the sulfateion as the example for illustration purposes). Hence when usingpotassium carbonate as the strip aqueous solution, the effectiveconcentration range is 0.1% to 50% K₂CO₃ at ambient temperatures; butpreferably 12-35% K₂CO₃ so that a concentrated potassium salt isproduced from S1-SO4 stage that is supersaturated with respect to K₂SO₄crystallization, making recovery of a solid K₂SO₄ product, with recycleof supernatant/filtrate preferred; and still more preferably 6-11% K₂CO₃so that the salt produced in stage S1-SO4 remains soluble so that aliquid product that is readily saleable as liquid fertilizer is made,and so to avoid the potential for hard scale to form in S1-SO4 unit whenthe product produced there is K₂SO₄, and most preferably 7.5 to 8.5%K₂CO₃ so that the maximum (about 10-11%) concentration of K₂SO₄solution, that does not crystallize at room temperature, is produced.

Similarly, when using sodium carbonate as the strip aqueous solutionbeing fed to the last strip stage, and optionally, although lessdesirably, to one or more of the others, the effective concentrationrange is 0.1% to 25% Na₂CO₃ (as anhydrous) at about 25° C. temperature,and greater at higher temperatures, but preferably 6-15% Na₂CO₃ so thata concentrated sodium salt of up to about 22% Na₂SO₄ is produced (asNa₂SO₄*10H₂O if T is <˜33° C., and anhydrous if T is >˜33° C.) fromS1-SO4 stage, and still more preferably about 10-15% Na₂CO₃ so that thesalt produced in stage S1-SO4 is soluble and sodium sulfate hydratesolids do not develop in S1-SO4 unit when the product produced there isNa₂SO₄ solution.

The SO₄ ⁼ and HCO₃ ⁻ ions so stripped by the carbonate ion using thesolutions defined above then accompany the aqueous phase from the laststrip stage (normally S4-SO4, but can also be S3-SO4, S5-SO4 or S6-SO4depending on the initial and residual (SO₄ ⁼) concentration objectivesof the feed and purified product waters of the extraction circuit), toS1-SO4 to become part of the final products as given above. In chemicalreaction form these reactions for stages S3-SO4 and SO4 are

{(R₄N⁺)₂SO₄ ⁼}_(E-phase)+M_(x)CO₃→M_(x)SO4+{(R₄N⁺)₂CO₃}_(E-phase)  (7)

and/or,

2{R₄N⁺HCO₃}_(E-phase)+M_(x)CO₃→{(R₄N⁺)₂CO₃}_(E-phase)M_(y)(HCO₃)_(z)  (8)

Similarly for optional stages S5-SO4 and S6-SO4 if they are used.Extractor Configuration to Maximize Separation of Trivalent Metal Ionsfrom Divalent Metal Ions

Using the example of acid mine drainage or acid rock drainage feedwater, the partially purified acid mine drainage aqueous steam generatedin extraction circuit E1 (operating at pH 3-5 for trivalent metals) isdirected to another extraction circuit (E2) for further processing at pH5-10 (for any residual trivalent and especially divalent metals presentin the water feed). At the lower pH, conversion of CO₃ ⁼ to CO₂ occurs,thereby requiring the R₄N⁺ to ion pair with any remaining anionic,especially sulfate ion in the case of acid mine drainage, not extractedin E1 stage, and to also form additional molecular metal ion complexclusters, colloids, flocs, particulates, and the like, containinghydrophobic quaternary moieties, as described above, with hydroxides andoxides of metal ions, mostly ferrous ion in the case of acid mine poolwater, collectively illustrated as {(R₄N)_(w)}{M(OH)_(n)}_(w) along withessentially separate and essentially soluble colloids of sulfate ionwith the quaternary ammonium ion extractant depicted as({(R₄N)₂SO₄}_(colloid))_(v). For these species “w” is the absolutenumber of charges per colloid, cluster, particle, and the like, and alsothe absolute number of charges on the sulfate ion colloids, to within azeta potential of 25 mV or less.

To separate the lower pH extracting metal ions from those extracting athigher pH, the extractant phase from above is then stripped with an acid(e.g. sulfuric acid or one of the other acids listed above) to removethe trivalent metal ions, and any more acidic +2 metal ions such ascopper(II), as metal sulfate concentrates, as described above, to thenyield {R₄N⁺HSO₄ ⁻}_(E-phase), depicted as {(R₄N⁺)₂SO₄}_(E-phase).Trivalent and certain divalent metal sulfates are recovered and areuseful as chemical specialties and commodities (see above). When thetrivalent metals present are ferric ion and aluminum ion, and also whenthe ferrous ion is co extracted and stripped with these ions, then aunique ferric aluminum ferrous sulfate product aqueous concentrate is soproduced and represents a unique composition of matter of value in largescale water purification applications that provide many beneficialadvantages over the separate components of ferric sulfate, ferroussulfate and alum. There are many other uses for this new composition ofmatter product (listed above).

One new composition of matter was discovered and produced using theinvention and consists of a blend of Al³⁺, Fe³⁺, HSO₄ ⁻, and SO₄ ⁼ ions,with a second formulation of these ions and Fe²⁺ ions at concentrationsand ratios that maximize municipal and industrial waste water treatment,municipal potable water, and/or industrial process water purificationprocesses. Conventional treatments involve the use of only one of thefollowing for these applications and at higher costs: ferric chloride,ferrous sulfate or alum. However, although these reagents are functionalin removing dissolved toxic metal ions from water and coagulatingsuspended solids and biologically derived biomass (from anaerobicdigesters), they all have serious shortcomings. First, though ferric ionis preferred for the above properties, the chloride content of ferricchloride promotes corrosive ion water distribution tankage and piping,leading to serious corrosion of copper plumbing and to Pb-based soldersin the water distribution system causing serious birth defect negativehealth effects. Hence ferric chloride coagulate has been reduced in usefor potable water production. A second problem with ferric chloride isthat ferric ion too rapidly hydrolyzes to FeOOH precipitate, thuspreventing it from dispersing well enough in the treated water to enableits efficient use in removing other pollutants of concern, especiallyarsenic As and phosphate ions (PO₄ ³⁻). Alum (an aluminum sulfatehydrate) is desirable in water treatment due to ease of dewatering butsuffers from the deficiencies of narrow operating pH range andespecially in forming excessively voluminous sludge, some ten times ormore that formed by ferric salts. Ferrous sulfate avoids the problemwith the use of chloride ion and is water soluble around neutral pHenabling it to be thoroughly dispersed so to more effectivelyprecipitate toxic contaminants such as As and PO₄ ³⁻. However, ferrousion is a poor coagulant and flocculent and must be air or chemicallyoxidized to provide coagulation and flocculation, a slow or costlyprocess, respectively.

The new composition of matter now provided by this invention avoids theabove problems while maintaining all of their benefits. The metalsulfate concentrate product is thus unique for purifying a broadspectrum of waters for many needs. It does not contain the corrosivechloride ion, it does contain ferric ion in abundance for fastcoagulation and flocculation, it does contain ferrous ion that providesa fully soluble form of iron to effectively precipitate As and PO₄ ³⁻,and lastly, it does contain a low amount of aluminum ion that issufficient to both perform coagulation and flocculation with a “filteraid” effect on the result and sludge filtration and dewatering steps,and yet the precipitate is dense due to the ferric content thatcollapses the otherwise voluminous AlOOH precipitate.

This new composition of matter is formulated Fe^(III) _(x)Fe^(II)_(y)Al_(z)(SO₄)_(w).H₂O and is provided in both aqueous solution andsolid forms. The equivalent ratios (dry conditions where q=Ø), based oncharge balance, are,

(X×3)+(y×2)+(z×3)=2  (9)

As an example, for x=1, y=1 and z=1, then the value of “w” is(3+2+3+)/2=w=4, provides the empirical formula ofFe^(III)Fe^(II)Al(SO₄)₄.qH₂O with a nonhydrous formula weight of 523g/mole. Another example, based on acid mine drainage water feed isx=0.1, y=0.8 and z=0.1, giving a value for w=(0.3+1.6+0.3)/2=1.1, orFe^(III) _(0.3) Fe^(II) _(1.6)Al_(0.3)(SO₄)_(1.1) for a formula weightof 159 g/mole. Specifically, this new material consists of the range ofratios: Fe^(III) _(0.03) Fe^(II) _(0.95)Al_(0.03)(SO₄)_(1.04) (FW of 155g/mole), to Fe^(III) _(0.95)Fe^(II) _(0.03)Al_(0.5)(SO₄)_(2.2) (FW of280 g/mole). The total composition of the aqueous solution of the newmaterial, Fe^(III) _(x)Fe^(II) _(y)Al_(z)(SO₄)_(w), produced by theprocess of the invention from acid mine drainage feed water is (dryweight) at least 0.1%, and preferably 1 to 5%, and most preferably 5 to20%. All percentages are by weight in aqueous solution. The solidmaterial is prepared by taking the metal sulfate concentrate from theprocess of the invention and either cooling it to 0 to 5° C. andallowing the material to crystallize, or by drying the material usingheat, evaporative cooling, or the like, until sufficient water has beenremoved to enable crystallization of at least a part of the originallydissolved, and preferably precipitating at least 90% of the dissolvedsolids, where upon a substantially amorphous, granular, air sensitivematerial is produced. Up to 50% of moisture is left in the solid toretain the materials reactivity and water solubility. Oven drying thematerial, especially in the air or other source of oxygen gas, is notrecommended since at least a portion of the material then would formamorphous mixed oxides of iron and aluminum, a portion would darken tomagnetite (mixed ferric and ferrous particulate) would be produced, andthe sulfate would be converted back to sulfuric acid making the materialhazardous to handle. This new and unique material is made by selecting astrip acid, specifically sulfuric acid in this case, that is used tostrip the Fe(II), Fe(III) and sulfate ion components of a metal-loadedextraction concentrate at an E/A ratio such to provide substantialconcentration enhancement, at least 10 fold, preferably 50 fold, andmore preferably 100 fold. In addition to these parameters the productconcentration of this unique material produced is determined also by thenumber of acid strip stages used and the number of internal recyclesachieved. The S1-M strip stage is the most influential on metal ionconcentrate. At the minimum, stripping must be sufficient to strip asubstantial portion of the metal ion content of the colloid loadedextractant phase while mixing for up to 30 min, preferably 15 min, andmost preferred up to 6 minutes. Normally more than 99% of the metal ionsof the colloidal floc are stripped into the mixed metal sulfateconcentrate due to the multi-stage counter-current liquid-liquidstripper design of preferably three contact stages.

An important and unique provision of this invention is that, in one unitoperation with a relatively small size it provides:

1) Removal of toxic cations, including the very slow reacting ferric,ferrous, nickel, and aluminum ions, and the like, even present at veryhigh levels, to levels compliant for surface water discharge, potablewater, and other regulated requirements,2) Neutralization of the acidity that characterizes many waste streams,including acid mine drainage, acid rock drainage, natural gas well, andthe like, discharge waters,3) Removal of the very difficult to remove sulfate ions and lowering ofsulfate-based salinity by 90 to 99% or more,4) Lowering of water total dissolved solids,5) Achieving the above separations simultaneously and quickly, therebyachieving a process with a relatively very small size footprint for highflow rate systems,6) Significantly, by means of this invention, highly toxic contaminantssuch as Pb, As, Se, Hg, Cd, and the like; and manyradioisotopes/radioactive contaminants such as Tc, I, and the like canbe effectively removed from the water,7) Purification even of waters contaminated with extremely problematic,complex mixtures of contamination ions in waters containingslowly-reacting ions, colloids, highly-toxic metal ions, and up tothousands of mg/L (ppm) of dissolved salts.

A particularly problematic example is that of mining, especially that ofcoal mining, where the mining activity has exposed gangue minerals leftin the mine and mine tailings to erosion by air, water, and microbialaction. Sulfidic minerals, such as pyrite, FeS₂, are commonly found inmany geological strata, and especially in reducing ores such as coalsand metal sulfide ore bodies. As described thoroughly in the literaturereview spanning many decades, coal, metal and other mining operations,and natural weathering fissures have allowed water, air and microbialaccess to these reducing substances. These conditions promote theoxidation of the sulfidic minerals to water soluble metal sulfates,especially ferrous sulfate solution, as shown by the series of chemicalequations below. In the specific case of iron, a complex series ofreactions can occur (J. Skousen, et al, Jun. 1, 1998, “Acid DrainageTechnology Initiative (ADTI),” Published by The National Mine LandReclamation Center located at West Virginia University in Morgantown, W.Va., Handbook of Technologies for Avoidance and Remediation of Acid MineDrainage) with the net total dissolved solids level and acidityincreasing result as follows (aided by microbial action):

2FeS₂+7O₂+2H₂O→2Fe²⁺+4SO₄ ²⁻+4H⁺  (10)

4Fe²⁺+O₂+4H⁺→4Fe³⁺+2H₂O  (11)

4Fe³⁺+12H₂O→4Fe(OH)₃*+12H⁺  (12)

FeS₂+14Fe³⁺+8H₂O→15Fe²⁺+2SO₄ ²⁻+16H⁺  (13)

4FeS₂+15O₂+14H₂O→4Fe(OH)₃*+8H₂SO₄  (14)

-   -   * solid at pH >2.5

These reactions produce an acidic ground water of primarily ferroussulfate, the main soluble metal ion product, containing very high levels(often above 1000 mg/L) of total dissolved solids, being largely acontribution of acid attack on dolomite and limestone carbonate-basedminerals (Reaction 15 and 16 respectively) which contribute Mg, Mn andCa, which also report to the water in sulfate form. Though bicarbonateion can also be present in some alkaline acid mine drainage waters,waters of pH less than about 6 will have converted most of the carbonateof these minerals into gaseous CO₂ and so little alkalinity will isfound.

2H₂SO₄+MgCa(CO₃)₂→2H₂O+2CO₂(g)+Ca²⁺+Mg²⁺+2SO₄ ⁼  (15)

-   -   From dolomite    -   Eq. 14

H₂SO₄+CaCO₃→H₂O+CO₂(g)+Ca²⁺+SO₄ ⁼  (16)

-   -   From limestone    -   Eq. 14

Hence, due to Reactions 12 and 14, acid mine drainage, acid rockdrainage, natural gas well brine, and the like, waters will tend to havefar more dissolved sulfate ion than dissolved iron, the pH will behigher than Reactions 10 to 14 would predict alone, and would be low inferric ion content. Hence, for the maximum benefit, the technology ofthis invention is best applied to the acid mine drainage water drawndirectly from the mine pool and not from the discharged surface watersin that much more of the metal ions can be recovered instead of lost toFeOOH precipitation in the affected natural stream or other water body.

These drainages and associated ground and surface waters contain one or,more often, many of the toxic or highly concentrated contaminants listedin Tables 2A, 2B, and 2C which provide an analysis of a representativesamples of acid mine drainage.

Conventional technologies have been found to be unsuitable for purifyingsuch waters as acid mine drainage, many mineral processing waters, gaswell brines, and the like. Typical concentrations of the contaminants ofconcern in such waters are several fold above the 1000 mg/L removable bysuch prior art technologies and the ferrous ion rapidly air oxidizes toFeOOH precipitate rapidly plug alternative technologies. The flow ratesof such ground waters, especially as they seep to the surface andcontaminate ground water for wells, streams and rivers far exceed thepractical treatment maximum of these alternatives.

-   -   Industrial processing of metals (surface finishing). These        plating and painting shop rinse waters contain one or more of        the toxic or highly concentrated contaminants listed in Tables        2A, 2B, and 2C and also toxics such as hexavalent chrome,        cyanide ion, lead, cadmium, nickel, copper, iron and others.    -   Discharges of agricultural water from concentrated animal        feeding operations, CAFOS, and the like. These drainages and        associated ground and surface waters contain one or more of the        toxic or highly concentrated contaminants listed in Tables 2A,        2B, and 2C, and more particularly also contain problematic        macronutrients, especially phosphorus (P) and nitrogen (N)        containing organic and inorganic compounds. P and N nutrients        are the bane of agricultural runoff waters because they        stimulate plant algal, and then bacterial, growth that far        exceed natural waters' ability to handle it, resulting in        deoxygenated zones (“anoxic” zones) in natural bodies of waters        extending from the farm, such as streams and creeks, leading to        ponds, lakes and rivers, and then to large lakes, like Lake        Erie, gulfs, like the Gulf of Mexico, and bays, like Pamlico        Sound (all in the USA). These dissolved oxygen “dead zones” are        lethal to aquatic fish and thereby render these streams and        agricultural practices problematic to the public recreational,        fishing industry, and government regulatory bodies.

Most phosphorus and nitrogen containing compounds in fresh manure andurine animal waste, or from animal feed decomposition processes takingplace in storage bins, occur as oxidized phosphate in the case of P butreduced amines in the case of N, the later some inorganic, primarilyammonium ion, and the rest as well known biochemical organo aminecompounds. It is also well known that the P component in agriculturalrunoff water is by far the most impactful micronutrient on algal blooms.Since the present invention readily extracts anions, like phosphatesthat are present in these problem waters, a substantial improvement inwater quality emanating from farms can be made using the invention.Likewise, after operating the residual organic and ammonical nitrogencomponents oxidatively using established physical chemical and/orbiological operations to convert the nitrogen to the nitrate anion, thenthe nitrogen component could also be so removed from the water andconcentrated for fertilizer use, for example as needed by internationalgrain grower farms. In this manner, agricultural waters can be purifiedto enable their problem-free discharge to surface waters without thenegative impacts described above due to the payload of nutrients. Mostpreferred is to first filter out any suspended solids by conventionalfiltration prior to treatment according to the present invention.

Method (Operation of Process)

Acid mine drainage/acid rock drainage feed water is used to illustrateapplication of the preferred operation of the invention. Coal mine acidmine drainage water (Tables 2A, 2B, and 2C) is typically mostcontaminated with the metal ions: (“M” Series) Fe(III), Al, Cu; (“N”Series) Fe(II), Ni, Co, Zn, Mn, Ca and Mg and highly contaminated insulfate anion. Sulfate ion is a notoriously difficult anion to removefrom water, and especially in a practical manner for the cases of unmetneed involving high to very high flow rates (e.g. 10 to 10,000 gal/min).Typical acid mine drainage water contamination and flow rate levels varybroadly by source, season, and physical characteristics of the aquiferinvolved. This invention enables the removal of most or all of thesecontaminants to levels low enough to allow widespread use of the waterno matter the variations in inlet contamination and flow rate levels.Contaminant levels can usually be reduced to below maximum permissibledischargeable levels in one treatment, and even routinely to levelsallowed by primary and secondary government drinking water standards.Therefore the product water is safe for aquatic life promoting tourisminvolving such river and stream systems. Even more, such purity levelswould enable the product water of the invention to be used for potablewater, feed for potable water production plants, farm water use, and/orfor industrial process water. As described below, the ability for thishigh level of purification is provided by the invention via a novelcombination of new process chemistry, newly designed simultaneousfloc/liquid/liquid contactor devices, and the newly discovered processoperations needed to deploy these technologies. Referring now to FIGS.1A and 1B, a detailed description of this integrated invention follows.

As used herein the term ionic ions species includes one or more of ananionic and/or a cationic ions.

A broadly applicable and detailed version of the invention is given inFIGS. 1A and 1B. This water purification process provides new and usefulprocess chemistry, water purification devices, and methods. The majorfeatures of the invention, the operation of which is detailed in theexample below, include a novel flocculation-liquid-liquid extractiondevice that possesses a unique design for floc and slurry handlinghardware to enable metal hydroxide/sulfate colloids and flocs to behandled in modified liquid-liquid contacting devices. Such flocs andslurries rapidly shutdown conventional and other prior art liquid-liquidextraction apparati. In one embodiment, this unique hardware consists ofa liquid-liquid contacting device fitted with a extractant phase flowcontrol consisting of gradually narrowing channeling gates feeding it toa smooth, preferably about 30 degrees sloped-surface, that feeds it to asmooth and broadly rounded over-flow type weir called a floc weir fittedwith a sloped discharge ramp of most preferably about 30 degrees fromthe vertical. These features enable the colloids but especially flocsand slurries produced by the unique process chemistry of the inventionto be separated from the water phase being purified in the process. Theunique separation process chemistry of the invention is normal verydifficult to process but is required to achieve very short waterresidence times in the equipment which in turn provide the enormousadvantage of purifying very large volumes (10 to at least 10,000gal/min) of water practically and at much lower cost than conventionaltechnologies.

Referring to FIGS. 1A and 1B the detailed description of the inventionis as follows. Acid mine drainage water 1110, containing anionic andcationic components, is fed by gravity or by pump 1120 through valve1125 and line 1130 to one or more floc/slurry liquid-liquid extractors1210 fitted with specially designed floc/slurry handling decanters asdescribed elsewhere in this application. The level of contamination andthe degree of water purification desired determines the number of suchextractors deployed. Normally the selection of the number of stages andflow rate ratios used is made by the level of purity desired for theproduct water, especially the residual sulfate ion and toxic metalresiduals needed. Most preferably the arrangement of such extractors iscounter-current, but can be preferably cross-current, or leastpreferably co-current. If more than one contactor is provided, they aremost preferably arranged in counter-current liquid-liquid extractionconfiguration. This water is contacted only for a short period of timein each extractor, 30 seconds to 30 minutes are effective, however,preferably only 30 to 200 seconds, and most preferably about 45-90seconds, and still me preferably 60 seconds, and still more preferably60 seconds, with extractant phase (defined elsewhere in thisapplication), supplied from tank 1010 via pump 1020 and valves 1030,1032 and 1034.

Water Purification Using Only One Extraction Operation

In the simplest case, where only mostly Fe(III), (including Fe(II)converted to Fe(III) during processing), Al, Cu, (“M” Series of metalions) and a substantial amount of sulfate ion needs to be removed fromthe water to accomplish the purification objective, then only the Low pHflotation liquid-liquid extractor process arrangement is used. In thiscase at least one flotation liquid-liquid extraction extractor set 1210,containing at least one mixing stage forming at least one waterimmiscible heterogeneous fluid consisting of liquid (at least two phasesconsisting of an aqueous and water immiscible extractant phase(E-phase), and also a floc and/or colloidal emulsion, and/or optionallya solid/liquid slurry. Also included is at least one settlingcompartment, optionally, but not preferably, fitted with “picket fence”fluid flow disrupter gates, and where the settler compartmentnecessarily contains a flow constrictor positioned to guide the topphase (organic emulsion containing colloidal floc and/or slurry) towardsthe specially designed floc/extractant phase overflow weir at the backof the settler. In this manner extraction of at least a portion of theabove contaminants from the acid mine drainage (or other) feed water isaccomplished using the extraction process chemistry described elsewherein this application.

The metal ion and sulfate ion “loaded” extractant phase from 1210 exitsthe settler as a water insoluble floc and/or slurry commingled withliquid extractant phase optionally, but most preferably, to a decanter.The construction of the decanter (for FIG. 1A is to be presumed to be acomponent of F-LLX Extractor 1210) is similar to the settler andincludes the floc handling overflow weirs described above. However, themost preferred operation of the Extraction 1210 settler and associateddecanter in that the E/A interface of the settler is set high in theextractor settler to promote exiting of the E-phase and its contents asquickly as possible to the decanter, and low in the decanter tofacilitate releasing the bulk of the water from the extractor settler asfast as possible but releasing the extract with floc slowly from thedecanter to maximize its dewatering. This short residence time isvaluable since the acid mine drainage water flow rate is by far thefastest flowing fluid in the system (at least 5 to 10 times any of theother flows) and therefore dictates the size of the hardware need tomake the purification.

The E-phase with floc is dewatered maximally in the decanter prior tometal product production. This further dewatering is achieved by settingthe E/A interface at a medium position in the decanter, normally in the1/3 to 2/3 range level of the total fluid depth of the decanter. Thisinterface positioning allows a very clean/sharp separation of the waterfrom the floc material in the decanter. It also allows the E-phase tothicken as it releases water of entrainment, chemically formed waterfrom the dehydration of metal hydroxide colloids, chemical displacementof hydration water by the modifier (normally oil soluble isodecanol orother oil soluble yet polar component of the E-phase, such as estersand/or alkyl phenols). Critically, the fluid dynamic design of thedecanter internals provide the necessary promotion of continuousthickened E-phase floc fluid flow in the decanter which, if such deviceswere not present, would result in settler and decanter plugging by heavyprecipitates causing stoppage of flow. However, with suchnewly-discovered devices, the E-phase emulsion thickens, but does notcollect at the interface (known as “CRUD” in the conventional technologyand prior art, resulting in the requirement for maintenance and evenshutdown/cleanout), but instead flows smoothly and continuously to theE-phase exit chamber 1040, which can be a line, a chute, or the like, asdescribed elsewhere in this application, thereby transferring it to themetal ion stripper 1220 operation.

The acid stripper operation 1220 consists of one or more, mostpreferably counter-currently arranged liquid-liquid contactors, whichcan be of conventional design, but which most preferably is designedcapable of accepting flow of floc/slurry impregnated E-phase to themixer compartment via 1040. In the case where more than one strippercontactor exists in counter-current arrangement, this modification isonly needed for the first mixing compartment as the floc or slurry israpidly dissolved in the first strip mixer operation.

In the strip mixer operation, the floc/slurry loaded E-phase iscontacted with aqueous acid solution 1215. This acid is delivered to thestripper contactor via pump 1217 and valve 1219. Most preferably, thestrip acid is continuously internally recycled within each striperliquid-liquid contactor used in order to achieve both efficient use ofthe acid, by consuming it as completely as possible, and themaximization metal ion concentration in the metal sulfate productsolution. The greater the metal ion concentration of the aqueous stripsolutions the more useful and valuable is the metal sulfate product andthe more cost effectively it can be processed into items of commerce.With the invention, the extracted metals (and later sulfate, see below)can be concentrated many factors, for example from 2× to 200,000 times,and often to saturation points for the metal sulfate, carbonate orhydroxide solids produced (see below). It will be appreciated that thelower the concentration of solute in the acid mine drainage feed water,the greater the concentration factors that are theoretically possible.For example, 2000 ppm sulfate ion present in feed water 1110 can beconcentrated by the invention to a sulfate solution concentrate inoperation 1370 of 200,000 ppm, representing a concentration factor of100×. However, 200 ppm sulfate ion present in feed water 1110 can alsobe concentrated by the invention to a sulfate solution concentrate of200,000 ppm, representing a concentration factor of 1000×. In thismanner, valuable metal ions, for example Co or Ni, that are present atsub-economic concentrations, for example 0.1-1 ppm, can be concentratedto 1000 ppm in operation 1220, a 1,000-10,000× concentration factor,making these metals now economically and practically available forcommercial use. The invention achieves these attractively highconcentration factors by using a combination of sharp separation processchemistry both for extraction and stripping, low E/A ratios in theextraction stages, and then high E/A ratios with aqueous internalrecycle in the strip stages, in the stripper stages 1220 and 1250. Theamount of recycle acidic metal sulfate is maximized and controlled byvalve 1240.

The aqueous acid, now carrying the metal extracted from 1210 andrepresenting metal sulfate concentrate exits the metal ion stripper 1220via line 1230. If the metal concentrate is concentrated sufficiently forharvesting and/or there is a need to drain aqueous phase from stripper1220 to prevent excessive aqueous phase volume accumulation in thestripper settler compartment, then this product metal sulfate aqueousconcentrate is harvested via valve 3-way valve 1240 to allow flow ofthis metal sulfate concentrate to optional oil/water separator 1270. Anyrecovered extractant phase from oil/water separator 1270 is returned tothe extractant phase exiting the metal ion stripper circuit 1060. Inthis manner, especially preferred to be a counter-current multi-stagedesign, the E-phase is rendered devoid of most of the metal ionsextracted above in the extraction operation and now proceeds to thesulfate stripping operation via line 1060. The metal sulfate concentrateexits optional o/w separator 1270 via line 1280 and proceeds to optionalsolid/liquid separator 1285, for example a crystallizer, in-line filter,or other solids/liquid separator, the liquid proceeds to M metal sulfateconcentrate through line 1290 to metal sulfate product collection vessel1292. The optionally crystallized product or filter cake is collectedvia line 1287 to receiving vessel 1289. The extractor and stripper canbe one contactor each or more than one. If more than one, then they aremost preferably configured counter-current or cross current. The conceptof counter-current, cross-current and co-current are already well knownto those skilled in the art of industrial scale liquid-liquidextraction. For the invention, one or two stages of flotationliquid-liquid extraction each, with decanter is most preferred when theobjective is to remove only ferric ion, aluminum, and/or copper withtwo, or preferably three, stages of metal ion stripping operated withinternal recycle in the mixer E/A range of 1/20 to 20/1, and with theaqueous acid feed fed most preferably counter-current or cross-currentto the strippers and while the aqueous phase from the third metalsulfate strip (S3M) stage is being sent to the second metal sulfatestrip (S2M) stage, then being sent to the first strip stage (S1M) asaqueous phase of metal ion sulfate concentrate from the first stage(S1M) is harvested. Decanters are not needed for the strip stage so longas the settler volumes are sufficient as is normally the case. Mostpreferred is that at least 15-30 min of metal sulfate stripper settlertime is supplied. Although a third phase can form in the S2M strippersettler, this can be alleviated by altering the composition of theextractant phase or most preferably by lowering the concentration of thestrip acid to about 25%, though 2%-30% acid concentrations areeffective.

After metal sulfate stripping operation, the extractant phase 1060, nowloaded with bisulfate ion and devoid of at least a portion of thetransition metal and light metal ion content, and preferably devoidof >90% of the transition metal and light metal ion content, and morepreferably devoid of >99% of the transition metal and light metal ioncontent, and most preferably devoid of >99.9% of the transition andlight metal ion contents, flows preferably by gravity through line 1080to the sulfate stripper operation 1370 where it is stripped of its anioncontent, sulfate ion in this example of acid mine drainage waterpurification, using one, or preferably 4 to 6 counter current,internally aqueous recycled, stages. Sulfate stripping is accomplishedin operation 1370 by contacting it counter-currently with an aqueoussolution of carbonate ion, hydroxide ion, a combination of the two,and/or bicarbonate solution, carbonate ion solution, or a combination ofthe two. Suitable cations for these anions are most preferably sodium,potassium or ammonium ions, or less preferably, lithium ion, or anycombination of these. Sodium cations and/or potassium solutions are themost preferred, i.e. Na₂CO₃ solution, and/or K₂CO₃, with or without KOHor NaOH. Although one mixer-settler is functional, preferably three tosix mixer-settlers arranged in counter-current flow is more preferred,and five is most preferred. The criteria for the number of stages to useare the amount of sulfate ion in the feed water, the E/A ration in thesulfate strip circuit, and the target low level of residual sulfate ionsought in the finished water. Any residual sulfate ion left on theE-phase after sulfate ion stripping operation is carried back into theextraction operation with the acid mine drainage feed water and hencemay reduce the sulfate ion removal efficiency of the extractionoperation and therefore the residual sulfate ion concentration left inthe purified water product.

For illustration, using the sodium cation case for the carbonate aqueousfeed 1310 to the sulfate ion stripper 1370, and the case of no NaOH feed1360, fed to the sulfate stripping operation 1370 via pumps 1320 and1350 via valves 1325 and 1345, the sulfate ion is removed from theextractant phase into the aqueous phase where it forms a sodium sulfatesalt concentrate at a ratio depending on what was fed via lines 1330 and1340 via pumps 1320 and 1350 from reagent feed tanks 1310 and/or 1360.During this operation, the extractant phase becomes loaded withcarbonate ion, bicarbonate ion, and/or hydroxide ion, depending on theE/A flow ratio in the strip circuit and the setting of valves 1325 and1345 and the speed of pumps 1320 and 1350. For removing the above listedacid mine drainage metals from the extractant phase, other than Mn, Ca,or Mg, preferably only carbonate ion, is used to strip the sulfate ionfrom the E-phase 1060 and/or 1080 (see below), and is fed from tank 1310with pump 1320 via valve 1325 and line 1330.

Note, that with bisulfate ion, HSO₄ ⁻ loaded onto the E-phase flowingfrom the M metal sulfate stripper as described above, the first stage ofsulfate stripping involves converting extracted bisulfate ion intoextracted sulfate dianion as follows,

2(R₄NHSO₄)+Na₂CO₃→(R₄N)₂SO₄+CO₂↑+H₂O+Na₂SO₄  (17)

Where, the second stage of sulfate stripping, and any additional stagessupplied, accomplishes essentially complete sulfate ion removal by ionexchange, i.e.

(R₄N)₂SO₄+Na₂CO₃→(R₄N)₂CO₃+Na₂SO₄  (18)

Where maintaining low E/A ratios in the stripper mixers enables thebuild up of very high sodium sulfate concentrations (e.g., 40,000 to650,000 ppm as SO₄ ⁼ where 150,000-250,000 ppm is preferred for the caseof a Na₂SO₄ product, at 10,000-100,000 ppm, and, preferably40,000-60,000 ppm, for the case of a K₂SO₄ product. Higherconcentrations are possible if preparations are made to harvest theresultant slurries of crystals formed.). Complete (efficient) usage ofthe carbonate is achieved. Notice that the humid 1371 CO₂(g) productproduced above can optionally be collected from the headspace of thesulfate strippers (especially from the first stage of stripping, S1-SO4,in a multi-stage operation). This CO₂ gas is a weak acid with naturalbuffering tendency for the pH range of about 6-9 and so, if needed, canbe used to bring the pH of the product water from pH >9 into this pHrange (see below) before discharging it to natural streams, used inagricultural operations, used as feed to potable water plants, used inindustrial operations, and the like. As will be shown below, high (>9)pH product water is produced only when >90% or the Mn, Ca and/or Mg areto be removed by the invention where pH values of the treated water canbe above 9, and as high as 12.

Once harvested, the high concentration of the sodium (or potassium)sulfate product stream exits line 1384 to optional O/W separator 1387through optional valve 1380. Valve 1380's main purpose is to preventpremature harvest (aqueous flow) to help insure full conversion ofbisulfate to sulfate dianion in S1-SO4 stage. Any E-phase captured bythe O/W is returned to stage S2-SO4 in the usual manner (O/W E-phaseflows are not shown in FIG. 2 to minimize clutter diagram, since theirrole and function are well known in the art, and because they normallycarry very little flow). The extractant phase return to theextractors/strippers is always to the next stage downstream of the stagewhere it was generated.

Exiting the O/W separator the sulfate concentrate flows to an optionalsolid/liquid separator 1390 via line 1389 if it is desired to removesodium sulfate crystals, potassium sulfate, or other solids such asmetal oxides and/or metal from the concentrate. Such crystals arecollected in product vessel 1392. The mother liquor then can becollected as liquid concentrate in tank 1398. Alternatively, bothproducts could be collected as slurry in tank 1398 by by-passing thesolids/liquid separator 1390, thereby obviating the need for line 1394and vessel 1392. After exiting the solid/liquid separator via line 1393,the sulfate concentrate flows through an optional trace E-phase sorbent,such as granular activated carbon (GAC) 1391 and/or other hydrophobicsorbent materials, then exits via line 1396 and is collected in tank1398.

As the sulfate concentrate has a very high ionic strength, little if anyE-phase contamination of it should occur. And then, any entrainedE-phase should have been captured by the O/W separator 1387. Hence theuse of (GAC) 1391 is not expected to be needed in most instances but ifused, then GAC use rate is minor and is used to remove any E-phase odorfrom the sulfate product.

The pH<5 raffinate exits extractor circuit 1210 via line 1420. Beforereleasing this water, it is most preferred to pass the water through O/Wseparator 1460 via valve 1430 and line 1450. The O/W separator effluent1470 can optionally be filtered and deodorized via solid/liquidseparator 1480 and GAC 1495 to produce clear, colorless and odorlesswater in line 1496 that is used immediately or collected in astorage/surge vessel 1494. The minor amount of particulates collected1492, normally of a rust-like character, via line 1491 are nontoxic andcan be discarded.

Water Purification Using Two Extraction Operations

In certain cases, depending on the requirement targets for the purifiedwater, it may be desirable to remove trace metal ion contaminants of themostly divalent metal ions, for example Ni, Co, Zn, Fe(II) as well asMn⁽²⁺⁾, (the “N” series of metal ions of FIG. 1A) and/or to achieve evengreater removal of sulfate ions. Additionally, it may be desired toseparate these metals from those collected in the pH<5 extractionoperation described above. In these cases two extraction operations,each with one or more liquid-liquid contactors, are recommended andprovided. This second operation is now described with the two variantsmentioned.

The first operation is as given above and results in an extractor mixerpH and raffinate pH of about 5. Adding a second extractor circuit 2030,which receives ˜pH 5 water from the first contactor circuit 1210 via thecombined settler and decanter aqueous overflow line 1440. The E-phase isfed to the second extraction operation from the common E-phase surgetank via line 1025 and pump 1020 and opening valve 1032. To operatecross-current, valve 1030 is open three way such than E-phasesimultaneously flows to both extraction operations 1210 and 2030.Alternatively, 1030 can be closed to 1210 and open to 2030, so that nowfirst extraction circuit 1210 receives extractant phase from secondextraction circuit 2030, i.e., a counter-current arrangement.

If ferrous ion is included in the acid mine drainage water feed 1110then it is normally present in large amounts relative to ferric ion.Hence extraction operation 2030 can again form a large amount of floc ofdivalent metal ions and so this contactor circuit is also of thespecially designed flotation liquid-liquid extraction type of theinvention, described in detail above, and is accompanied by a flotationliquid-liquid extraction decanter, also of the invention and previouslydescribed.

The operation of the second unit is as follows. Depending on the E/Aratio, number of contactors per circuit, the concentration of extractantin the extractant, and/or the number of equivalents of the above metalspresent, the final pH of the extraction can be about 7.5 to 9. Thiscondition efficiently removes divalent metals which can be onlypartially removed by the first pH ˜5 extraction. The attainment of pH 9helps Mn removal to about 90% and this higher pH can be promoted byincluding some hydroxide ion on the extractant that can be provided fromsupply tank 1360 via pump 1350 and valve 1345 through line 1340. It isone important aspect of this invention that basicity is brought into theextraction process without the accompaniment of water soluble cations.This accomplishment is made possible by using the water insolublecationic quaternary ammonium ion salt of basic anions. This featureenables the substantial decrease of the total dissolved solids of theproduct water relative to the feed water.

The pH 7.5-9 raffinate, now clear of substantial contamination canrepresent very desirable product water. Before releasing this water, itis optionally, but most desirable, to pass the water through O/Wseparator 2090 via optional valve 2070 and line 2080. (Valve 2070 isonly needed when additional extraction stages are added as describedbelow). The O/W separator effluent 2110 can optionally be filtered anddeodorized via solid/liquid separator 2120 and GAC 2125 to produce theclear, colorless and odorless product water in line 2140 that is usedimmediately or collected in a storage/surge vessel 2150. The minoramount of particulates collected 2130 via line 2115 are nontoxic and canbe discarded. Due to the scale of some of the acid mine drainage waterstreams, the filters used could be sand filters, clarifiers, simplein-line filters such as plate and frame, drum or belt filters,centrifuges, and the like. The GAC is again for odor control of clearwater product and so is optional and represents a small use rate.

The metal and sulfate ion loaded extractant phase can be process eitherof two ways according to the invention (for convenience the two optionswill be referred to as options AA and AB). These can be operatedsimultaneously, alternated with acid mine drainage feed content,economic conditions for the products produced, and so on.

In option AA, the loaded extractant is merely sent forward to the firstextraction phase via line 2040 and where valve 2042 is open to flow tothe pH 5 Extraction operation but closed to line 2044. In this mode,valve 1030 is optionally closed so that the entire extractant phase flowis through valve 1032 to the pH 7.5-9 Extraction operation. In thismanner all the metals extracted report to the metal sulfate extractorforming a combined “M+N”-Sulfate concentrate 1292 in the identicalfashion described above for the M-Sulfate product. After stripping ofthe metals, the sulfate is stripped in a fashion also identical to thatdescribed above previously.

Option AB differs in that it provides separation of M-sulfate andN-sulfate products using cross-flow contacting of E-phase with the acidmine drainage water and stripping the two loaded phases separatelyforming two concentrate products, M-SO4 and N-SO4. In this case pH 5Extraction operation is performed exactly as described above includingreceiving E-phase feed via valve 1030 and producing M-SO4 product 1292.However, the N-loaded E-phase formed in 2030 is sent to a separate acidstripper 4000 via line 2044 by diverting the flow from 2030 using valve2042. In this case a N-Sulfate concentrate is produced in a fashionidentical to M-Sulfate and using equivalent process hardware, includinginternally recycle aqueous via line 4020 and with the strippedextractant phase flowing forward to sulfate stripping 1370 via lines1070 and 1080. Hence the N-Sulfate concentrate, when harvested, is sentvia line 4010 and valve 4030 (which again can just be an overflow weir),through an optional O/W separator 4040 to an optional solid/liquidcrystallizer 4060 via line 4050 if solid product crystals are to beseparated from liquid product concentration (not preferred). The solidcrystals can be collected in vessel 4070 via line 4080 and the liquidconcentrate in tank 4090 via line 4095. Most preferred is just tocollect one product slurry into tank 4090. Deodorizing is again optional(and not normally needed or preferred), for example by GAC treatment.

Whether option AA or AB is preferred can depend upon the level of Fe(II)in the acid mine drainage water relative to Fe(III). If a high level ofFe(II) exists (this is often the case when acid mine drainage feed wateris derived directly from wells or abandon coal mine shafts), then thisFe(II) will be extracted in pH 7.5-9 Extraction operation and thereforethe more valuable but dilute “N” metals, i.e. Ni, Co, Zn and Mn, will becollected with the large volume of Fe(II)-sulfate concentrate produced(if air oxidation of Fe(II)→Fe(III) is prevented by limiting air accessto the extractor mixers, this is easily accomplished by simple tankcovers and by not sucking air into the feed lines to the mixers). Henceif there is very little Fe(III), Al and Cu present (“M” metals) thenthere is no motivation for operating two separate stripper circuits andcollecting all the metal sulfates in one product as a “concentrate” thatcan be further refined into saleable products in a side and much smalleroperation is preferred. Hence the acid mine drainage feed watercomposition determines the most preferred mode of operation.

It is noted that in the above pH 7.5-9 operation, the Mn is only about90% removed. Hence, depending on the fate of the product water(discharge or feed to potable water production plant, etc.) it may bedesirable to add a third water purification operation (see below).

Water Purification Using Three Extraction Operations

The need can arise for a third purification extraction operation in somesituations where the water is to be used rather than discharged. In acidmine drainage this arises where the intended use for the purified wateris for feed to a potable water production plant, and industrialoperation, and the like, where hardness metal ions, still lower TDS,lower alkalinity and/or less Mn contamination levels reduction isdesired.

For example, to reduce Mn levels to below secondary drinking waterstandards of 0.05 ppm for the case where the acid mine drainage feedwater level is sufficiently high that this Mn residual level is notattained in the raffinate from the first two extraction operations 1210and 2030 despite removing 90% of it. In this case, the pH of the waterneeds to be adjusted to higher pH to convert the Mn to MnCa(CO₃)₂particulate that can be filtered out. As the acid mine drainage feedwater is already far excessively contaminated in total dissolved solids,e.g. several to ten times the level permissible for drinking water. Itis needed to make this pH increase adjustment without the addition orformation of additional salts in the water. The invention provides thisunexpected capability as will be described below.

In addition, if the acid mine drainage feed water was initially veryhigh in sulfate ion concentration, there may be a need to remove thesulfate ion concentration to still lower levels to enable use, ratherthan discharge, of the water.

Due to additional floating floc formation involving the Mn, Ca and/orthe residual sulfate ion, the third extraction step is also mostpreferably performed as another flotation liquid-liquid extractionoperation (see above for description of the first two flotationliquid-liquid extraction operations). The preferred configuration forthese additional two steps is the flotation liquid-liquid extractionoperation but without a decanter. Decanters are no longer needed sincefloc formation is substantially reduced in the subsequent contactsrelative to the first two extractions. However, instead, particulatesform and the low ionic strength and higher pH of the water product aftera third extraction operation still higher pH results in at least aportion of the contacted fluid exiting the mixer compartment of 3020 vialine 3050 and valve 3060 to be a emulsion that requires additionalprocessing to separate (FIG. 1A and FIG. 1B)). This unstable emulsion isan ionic colloidal complex between the MnCa(CO₃)₂ solid particulates andthe Aliquat-SO₄ ⁼ species. Hence a solid/liquid (S/L) and liquid/liquidseparation is needed. Such separator can be a filter, a semi-continuouscentrifuge, hydrocyclone or other solid/liquid (S/L) separating device.The extractant phase formulation can also be adjusted to encourageseparation of these emulsion components for example by increasingmodifier concentration change diluent, blend diluents, and the like.

Freshly prepared extractant phase can also produce a “milk” coloredaqueous phase that is believed to arise from hydrophobic tertiary aminecomplexes that form from E-phase manufacturer impurities at the higherpH of the third extraction due to deprotonation of the amine. This milkdoes not contain significant amounts of Mn and can be discarded. Theextractant can also be purified of tertiary amine impurities prior tousing the extractant in the first place by acid washing, washing withcopper sulfate solution, and the like. How these process features areperformed is described below.

This third extraction operation is accomplished by sending thepretreated water from the second purification operation described above(using either options AA and/or AB) via valve 2070 to line 2060 toflotation liquid-liquid extraction (or conventional LLX) contactor 3020.

The pH of the third extraction is in the 8.5 to 10.5 range. This higherpH is accomplished by including hydroxide ion in the extractantregeneration operation of sulfate stripping. Most preferably, NaOH isused since it forms the far more soluble Na₂SO₄ product 1398, but KOHcan also be used if more dilute sulfate product is acceptable. Thehydroxide is fed to the sulfate stripper from tank 1360 via pump 1350and valve 1345 through line 1340.

At the higher pH, the concentration of carbonate ion becomes significantcausing the particulates of MnCO₃ to form; with any Ca present formingCaCO₃. Mg is not removed at these conditions requiring higher pH (seebelow). Particulate formation is desirable as it allows the removal ofthese Mn and Ca contaminants with a corresponding decrease in totaldissolved solids since the OH— can be introduced from via the cationicE-phase. The overall reaction is,

Mn²⁺+2HCO₃ ⁻+SO₄⁼+Ca²⁺+2(R₄N⁺OH⁻)_(E-phase)→MnCO_(3 particulate)+CaCO_(3 particulate)+((R₄N)₂SO₄)_(E-phase)+purifiedwater_(pH 8.5-10.5)

This product mixture forms an extractant with floc and an aqueous phaselow in Mn, Ca and SO₄ ⁼ content that exits contactor 3020 via line 3050and valve mixture exits the O/W separator 3040 via line 3065 and entersthe solid/liquid separator 3070. In the solid/liquid separator 3070, theMn(CO)₃ and Ca(CO)₃ solids 3090 are separated from the mixture via line3080. The extractant phase is sent via line 3075 to the secondextraction operation 2030 for additional stripping and regeneration.Although most of the E-phase separates from most of the aqueous phase,the low ionic strength of the product mixture, and the high ioniccharacter of the quat:sulfate ion cluster results in some formation of amicro emulsion milky white product water. Analysis of this microemulsion shows it to be low in Mn and SO₄ ⁼. The micro emulsion isbelieved to be caused of surface active impurities in the Aliquat 134,and these are known to wash out from the system over time, and so is anon-issue.

The extractant phase floc, a brown emulsion, contains the extracted Mn,Ca and sulfate ion. This floc is dewatered and thickened in the decanteras was done in the first two extraction operations. The E-phase flocthen could be stripped of metals using acid as before, filtered,centrifuged, or other solid/liquid separation 3070 to gather the Mn andCa in particulate form 3090 via line 3080. The extractant 3075 is thensent to sulfate stripping directly (preferred) or to the secondextraction operation, or to a fourth extraction operation (see below).The stability of this emulsion could also be reduced using a weaker(particulate) or stronger ion pair solvating E-phase formulation.

The pH 8.5-10.5 raffinate exits the solid/liquid separator 3070 via line3075 and valve 3100. The pH of the purified water can be reduced to thepH 6-9 range by sparging inline 3105 with the CO₂ gas 1371 via line 1372and valve 6000. In addition, clarified raffinate 3110 can be sent tocontractor 5000 for Mg removal. The pH 6-9 raffinate exits valve 3100via line 3105 and enters GAC 3120 for deodorization. The deodorized andpurified water exists GAC 3120 via line 3125 is collected in a storagevessel 3030.

Water Purification Using Four Extraction Operations

The purified water from the above-described four extraction stages isfairly pure. However, if Mg levels are significant then dissolved Mgsulfate and carbonate salts still would be present in the water. In thecase of acid mine drainage this can represent hardness levels in therange of about 400-900 ppm, too high for certain end use applications.For example for the case of acid mine drainage water with 140 ppm and1600 ppm SO₄ ⁼, 130 ppm Mg and 135 ppm sulfate can still exist in thewater giving rise to high total dissolved solids residuals.

Although Mg is not removed efficiently in any of the above-describedflotation liquid-liquid extraction operations, it was discovered that Mgcan be removed without employing conventional water softening by using aflotation liquid-liquid extraction operation 5000 at still higher pHwhere the hydroxide ion is again introduced by the E-phase via valve1034. It was discovered that in a flotation liquid-liquid extractionoperation, hydroxide ion can be effectively brought in with theextractant phase (added at the 1370 via 1360) where the pH is raised to10.5-12, preferably 11-11.5. In addition, clarified raffinate 3110 canbe sent to contractor 5000 for Mg removal. The process chemistry isbelieved to be as follows:

Mg²⁺+SO₄⁼+2(R₄N⁺OH⁻)_(E-phase)→Mg(OH)_(2 particulate)+((R₄N)₂SO₄)_(E-phase)+purifiedwater_(pH 10.5-12)

An emulsion comprised of Mg(OH)₂ particulates, aqueous phase, andextractant phase exits contractor 5000 via line 5010 and valve 5015 toline 5020 where the emulsion enters the optional O/W separator 5030. Theemulsion exits O/W separator 5030 via line 5035 and enters thesolid/liquid separator 5040. In the solid/liquid separator 5040, theMg(OH)₂ solids 5060 are separated from the emulsion via line 5065. Theextractant phase is sent via line 5050 to the third extraction operation3020 for additional stripping and regeneration. The pH 10.5-12 raffinateexits the solid/liquid separator 5040 via line 5070. The pH of thepurified water can be reduced to the pH 6-9 range by sparging inline5070 via valve 5075 with the CO₂ gas 1371 via line 1372 and valve 6000.The pH 6-9 raffinate exits valve 5075 via line 5080 and enters GAC 5085for deodorization. The deodorized and purified water exits GAC 5085 vialine 5090 and is collected in a storage vessel 5100. The product wateris very pure, now being depleted of both M and N metals, Mn, Ca and Mg,and with a SO₄ ⁼ residual <20 ppm. This product water is of sufficientquality as feed for most potable, industrial, and agriculturalapplications.

Table 1 lists examples of the water contaminants removed simultaneouslyby the invention. Such contaminated water streams are often so highlycontaminated that no one technology is effective, or the contactingelement becomes fouled with other components of the mixture, or thewater is not sufficiently purified for release or reuse and must betreated again. Such waters may contain many contaminants of concernunder EPA regulations. Such waters often include sulfate ion, a severeproblem, in concentrations of 150-2500 ppm and even higher.

TABLE 1 Typical water contaminants that can be removed by the inventionClass of Contaminants Typical Contaminant Species Removed ANIONS Sulfateand bisulfate, selenate, tellurate, phosphates and hydrogen phosphates,organophosphates, organophosphonates, organophosphinics, polyphosphate(esp. ATP, ADP and AMP), arsenic (especially arsenate, organo arsenates,and arsenite), chloride, bromide, iodide, and pseudohalogen ions.Nitrate, nitrite, cyanide ion, sulfide ion (H₂S in equilibrium with HS⁻,H₂S and S⁼), oxometal ionic complexes including single ion and molecularclusters and colloids of molybdates (based on MoO₄ ²⁻), tungstates(based on WO₄ ²⁻), vanadates (based on VO₃ ²⁻), chromates (based onHCrO₄ ⁻, Cr₂O₇ ⁼, CrO₄ ²⁻), and the like, including their protonatedweak acid species, and particulate ion. CATIONS Cations capable offorming oxide or hydroxide ion colloids and precipitates (“oxohydroxoclusters or colloids”) including nickel, copper, chromium(III), ferric,ferrous, aluminum, manganese, cobalt, cadmium, zinc, Pb, Hg, Cd, and thelike; and many radioisotopes/radioactive contaminants such as U, Th, Pu,and the like hardness metal ions (Ca, Mg), and the like contaminantssuch as Tc. ORGANICS Oil-soluble organics, natural organic matter andthe like.Below are analyzed samples of acid mine drainage from sites inPennsylvania (Tables 2A and 2B) and Ohio (Table 2C).

The Tables show chemical analysis of a actual representativePennsylvania acid mine drainage water samples to illustrate the extremecontamination level that is not practically treated by any other knowntechnology. The stream in Table 2B is known to flow at about 10,000gal/min or more year after year resulting in many thousands of tons ofmetals and sulfate contamination of the environment. It is unsuitablefor potable water production, for industrial process water use, foragricultural use, or for providing aqua-tourism, or to support mountainstream life.

TABLE 2A Analysis of Sugar Camp, PA Acid Mine Drainage Water4841218/4841219 AMD Sample Analysis Name Units Result MDL Aluminum mg/l2.73 0.0802 Calcium mg/l 128 0.104 Iron mg/l 204 0.0522 Magnesium mg/l135 0.0135 Potassium mg/l 5.41 0.0503 Sodium mg/l 2.85 0.433 Cobalt mg/l0.873 0.0021 Manganese mg/l 71.5 0.0036 Zinc mg/l 1.26 0.0081 pH Std.Units 3.2 0.01 Alkalinity to pH 8.3 mg/l as CaCO3 N.D. 0.46 Alkalinityto pH 4.5 mg/l as CaCO3 N.D. 0.46 Total Dissolved mg/l 2050 38.8 SolidsSulfate mg/l 1620 60.0 Specific umhos/cm 2440 1.7 Conductance Acidity topH 3.7 mg/l as CaCO3 83.8 2.0 Acidity to pH 8.3 mg/l as CaCO3 413 2.0

TABLE 2B St. Michael, PA Acid Mine Drainage Water Analysis St. MichaelsAMD Feed Water 51944-01-18 Analysis Name Units Result DL Aluminum mg/l4.16 0.0802 Calcium mg/l 158. 0.0632 Iron mg/l 169. 0.0522 Magnesiummg/l 160. 0.0135 Potassium mg/l 5.56 0.0503 Sodium mg/l 2.50 0.433Cobalt mg/l 1.03 0.0021 Manganese mg/l 77.5 0.0042 Zinc mg/l 1.69 0.0081pH Std. Units 2.9 0.010 Alkalinity to pH 8.3 mg/l as CaCO3 N.D. 0.46Alkalinity to pH 4.5 mg/l as CaCO3 N.D. 0.46 Total Dissolved Solids mg/l2,310. 38.8 Sulfate mg/l 1,760. 60.0 Specific Conductance umhos/cm2,720. 1.7 Acidity to pH 3.7 mg/l as CaCO3 N.D. 10.0 Acidity to pH 8.3mg/l as CaCO3 481. 10.0

TABLE 2C Rush Creek, OH High Al Acid Mine Drainage Water Analysis HighAluminum AMD Feed Water 51944-24-06 Analysis Name Units Result MDLAluminum mg/l 133. 0.0802 Calcium mg/l 398. 0.0632 Iron mg/l 1,130.0.261 Magnesium mg/l 151. 0.0135 Potassium mg/l 37.1 0.0503 Sodium mg/l23.2 0.433 Cobalt mg/l 0.0758 0.0021 Manganese mg/l 28.7 0.0042 Zincmg/l 0.970 0.0081 pH Std. Units 3.0 0.010 Alkalinity to pH 8.3 mg/l asCaCO3 N.D. 0.46 Alkalinity to pH 4.5 mg/l as CaCO3 N.D. 0.46 TotalDissolved Solids mg/l 7,350. 77.6 Sulfate mg/l 5,200. 150. SpecificConductance umhos/cm 5,560. 1.7 Acidity to pH 3.7 mg/l as CaCO3 N.D.20.0 Acidity to pH 8.3 mg/l as CaCO3 2,290. 20.0 N.D. = not detected

The invention is preferably used to treat the above mentioned waters andthe like before and/or instead of conventional waste water treatment toobtain the maximum beneficial impact by

-   -   Preventing the contamination of the environment and/or    -   Pre-treating contaminated waters before these waters are treated        in a manner that generates large solid and/or liquid wastes        thereby reducing the total amount of solid and/or liquid waste        produced.

This invention is particularly useful because it provides a broadspectrum water purification capability for a wide range of water flowrates, preferably flow rates of less than 1 to more than 10,000 gal/min.

Such extractant/co-extractant system is used alone or in combinationwith one or more modifiers to improve extractant solubility in thediluent, and/or with one or more water-immiscible diluents. Specificexamples are given in Table 3.

Suitable modifiers are water-immiscible terminal aliphatic alcohols ormixtures thereof. Preferred diluents are alcohols that are classifiednonflammable (flash point >140° F.), nonhalogenated, low-odor,aliphatic, either linear or branched, with a carbon number of 8-16, mostpreferably 9-13, or mixtures thereof. Specific examples are given inTable 3.

Suitable diluents can be water-immiscible aliphatic, aromatic solventsor blends of such solvents. Most preferred are solvents that areclassified nonflammable (flash point >140° F.), nonhalogenated, low-odoraliphatic, aromatic, or a blend of aliphatic and aromatic solvents. Thealiphatic diluent(s) can be linear but are preferably branched. Thearomatic diluent(s) can be unsubstituted aromatic liquids but arepreferably aliphatically-substituted aromatic liquid compounds.Extractant mixtures suitable for the invention contain at least 25%diluent (v/v), preferably 60% (v/v), and most preferably 85% (v/v).Specific examples of suitable diluents are shown in Table 3.

For a further embodiment of the invention suitable reagents forstripping of metal ions, regeneration, solute concentration, and/or pHadjustment include mineral acids that do not cause decomposition of theextractant. Suitable strip acids are preferably selected fromhydrochloric, sulfuric, phosphoric, blends of these and the like. Acidstrength is 2-50% in the feed solution and the S2M mixture is preferablymaintained at ˜pH 1-2 or less. S1 mixer pH is preferably maintained atpH 2-4, or less. Sulfate ions are subsequently stripped by basicreagents that do not cause decomposition of the extractant preferablyselected from alkali and alkali metal hydroxides, and/or alkalicarbonates. Base strength for sulfate ion stripping is preferably ˜pH10-11 or greater. Specific examples of suitable of reagents forstripping and regenerating, and pH adjustment of the extractant arelisted elsewhere herein.

TABLE 3 Typical Compounds Useful for the Extractant Phase Chemical Class(used alone or in Extractant combination Formulation with any otherLevel of Component extractant) Specific Compounds Preference ExtractantQuaternary N-methyl tri-(n- Most preferred Amines octyl)ammonium ionN-methyl tri-(n- Most preferred decyl)ammonium ion N-methyl tri-(n- Mostpreferred dodecyl)ammonium ion Aliquat ® 134 Most preferred Aliquat ®336 Most preferred Tri-octyl methylammonium ions Most preferred Mixtureof tridecyl- and trioctyl- Most preferred methylammonium ions HOE S 2706Most preferred Adogen ® 464 Most preferred Tri(C₈-C₁₀) methylammoniumions Most preferred R¹R²R³N⁺CH₃ R¹═R²═R³═CH₃(CH₂)₉— Most preferredR¹═R²═R³═CH₃(CH₂)₇— Most preferred and Most preferredR¹═R²═R³═CH₃(CH₂)₉— Blends of the above quaternary Most preferredammonium ions in any proportion Mono e.g. LIX-79 ® Most preferredguanadinium Quaternary N-methyl tri(n-octyl) phosphonium Functionalphosphonium Modifier b-isodecanol Most preferred (or decyl alcohol, orExxal ® 10) Isotridecanol Most preferred (or b- Tridecyl alcohol Nonylphenol aromatic Functional Dodecyl aromatic Functional phenol DiluentAromatic ® 150 Functional Aromatic ® 200 Functional Calumet ® MostPreferred 400-500 Conoco ® 170 Preferred Isopar ® M Preferred

Referring to Table 3, effective “Type extractant phase” extractcompounds of the invention. All of the components of the E-phase are oilsoluble with a total carbon number of at least eight (8), can be chargedor neutral, and can have additional functional groups such a halogens,ether linkages, ester linkages, alkyl phenolic, and the like, so long asthe extraction chemistry and the oil solubility of the reagent is notadversely effected relative to the descriptions herein.

Effective extractant compounds of the invention are all oil soluble witha total carbon number of at least eight (8), but preferably about 16,and most preferably about 25 or more, and can have additional functionalgroups such a halogens, ether linkages, ester linkages, aromatic groups,be linear or branched, blends of these, and the like, so long as theextraction chemistry and the oil solubility of the reagent is notadversely effected.

Referring now to FIG. 1A and FIG. 1B, a general overview of the processis shown. The figure shows extractant phase continuous recycle andoperation at pH up to about 11.0-12.0 (Mn and Mg recovery). Thuscontaminated water (e.g. acid mine drainage) is treated with extractantphase and within a short time water, together with solid products, ismoved to a solid/liquid separation apparatus where metal oxide/hydroxidecarbonate products {e.g. MnCO₃(CaCO₃), Mg(OH)₂} are removed. The waterthen flows to an oil/water separator that provides a purified waterproduct. Other contaminants are removed with LLX extraction in a loadedextractant phase that produces a floc. In addition to LLX extraction,product recovery and extractant phase regeneration, stripping is used toremove other metal oxide/hydroxide/carbonate products (e.g. Fe^(II),Fe^(III), Al, Ni, Co, and Zn) that were separated in the additional flocusing H₂SO₄ and (Na/K)₂CO₃, typically enables these metal sulfateproducts to be recovered.

Alternate liquid-liquid contacting apparatus can be provided asmixer-settler, columns, in-line mixers, contacting centrifuges, and thelike. Referring now to FIG. 3, one embodiment comprises four (4)extraction stages for metal cation and sulfate anion co-extraction, two(2) or four (4) acid stripping stages for metal stripping (depending onwhether metals are to be separated during production of metalconcentrates), and four (4) stripping stages for sulfate stripping andextractant regeneration. The overall process flow diagram using acombination of conventional and uniquely designed mixer-settlers isshown in FIG. 3.

The flow configuration of the device of FIG. 3 enabled the steady flowof heavy floc that was produced in E1 settler 316 to flow to the S1-M351 (shown as S1-MSO₄) stripper, where the floc is converted back to atwo phase liquid form manageable by conventional LLX hardware. The flocflow in E1 312, E2 313, and the associated decanter 314 is accomplishedby the “T” shaped configuration along with the floc flow handling weirscontained in the settlers and at the overflow point. These weirs aredescribed elsewhere in this application. Note that with the “T” design,the E-phase overflow from E2 settler 317, representing 325 b, the moreflowable floc shears off the heavy floc exiting from E1 settler 316, andhence insures that the floc from E1 settler 316 flows steadily. TheE-phase flow rates to E1 and E2 are adjustable and are used to optimizefloc flow continuity, extraction yield, and reagent consumptionminimization.

Also FIG. 3 illustrates the wraparound (space saving) means to increasethe residence time in the settlers of E3 341 and E4 342. Enhancedresidence time in these settlers is desirable as the total dissolvedsolids level in the purified water decreases and becomes slower to phaseseparate. As before, all of the strippers of the “T” configuration ofFIG. 3 are preferably operated counter-current with most preferredinternal recycle of aqueous. Note that the extractors were operatedcross-current with respect to the E-phase flow, and where the E3 341 andE4 342 stages are operated counter-current. Each loaded E-phase couldhave been stripped separately, but they are shown recombined in FIG. 3.Keeping these three loaded E-phases separate during stripping wouldallow production of three separate metal ion products (see DetailedDescription of the Invention).

Referring again to FIG. 3, the disclosed process uses a flocliquid-liquid extraction system to extract metal ions and sulfate ions.This typically is useful for cleaning up a contaminated andenvironmentally harmful acid mine water discharge stream. Due to theunique nature of the acid mine drainage stream and the floc-basedprocess chemistry, the extraction system apparatus has designrequirements that differ from conventional liquid-liquid mixer-settlersystems.

Conventional mixer-settler based LLX systems have the advantage ofself-regulating and low maintenance and labor cost. However, theconventional mixer-settler system performs poorly when dealing withmetal oxide colloids and/or floc, and would easily completely fail inminutes with the acid mine drainage feed water discussed here. A drivingforce for one aspect of the present invention is to change theconventional LLX contactor design to create a new way of dealing withthe metal oxides/hydroxide colloids, flocs, particulates or slurry whileretaining the advantages of conventional mixer-settler systems. This newdesign handles flocs along with liquid-liquid processing, or flotationliquid-liquid extraction.

Referring again to FIG. 3 that illustrates one embodiment according tothe invention as applied to acid mine drainage water cleanup using thesimultaneous extraction process of the invention in a specially designedapparatus 300. This apparatus 300 typically has a acid mine drainage(AMD) water inlet 301 and various interconnections for liquid flowreferred to as internal recycle lines 303, aqueous lines 304, andextractant lines 305. The water 301 flows into a T shaped extractiondecanter apparatus 311 made up of a first extractor E1 312, a secondextractor E2 313, and a dual decanter 314. The extractor E1 312 has amixer 323 which receives the incoming water 301 via an aqueous line.Here the water 301 is mixed with extractant solution 302 (the extractantsolution is loaded with an anionic base—in this case sodium carbonate)from the extractant solution storage tank 307 via an extractant line andpump 318. The mixture flows into the settler section 316 of extractor E1312 where the extractant 302 and water 301 interact to form a floc. Thefloc is allowed to separate typically by being allowed to float to thesurface, a flow guide 324 a helps the floc to flow over a floc weir 325a. The floc containing extractant, some water and captured ions flowsover the floc weir 325 a into the dual decanter 314. The aqueous portionof the material left behind in E1 settler 316 flows via an aqueous lineto the mixer 331 of extractor E2 313. Additional extractant solution 302from tank 307 flows via an extractant line and pump 319 to mixer 331where the water from E1 settler 316 is mixed with the additionalextractant solution 302 which then flows into the settler section 317 ofextractor E2 313. After also being allowed to form a floc and separatefrom most of the aqueous the floc is guided by flow guide 324 b over thefloc weir 325 b and into the dual decanter 314. The flocs from E1settler 316 and E2 settler 317 typically mix and are allowed to separatefrom entrained aqueous. The newly formed and/or reformed floc in thedual decanter 314 is guided by a flow guide 324 c to flow over the flocweir 325 c into a collector where the floc flows to the first stripperS1-MS04 351.

Aqueous depleted in ions in dual decanter 314 is typically withdrawnfrom the bottom and flows via aqueous line 330 to extractor E3 341 whereit enters at mixer 332, here extractant from an extractant line 331 fromextractor E4 342 also enters the mixer 332. After mixing the aqueous andextractant phases, the mixture enters the settler 333 of extractor E3341. The mixture is allowed to separate into an aqueous phase and anextractant phase. The extractant phase is withdrawn by an overflow weirinto a collector 334 having an outlet 335 and then flows via anextractant line into the mixer 352 of S1-MSO4 stripper 351. Aqueousflows to outlet 336 and via an aqueous line to the mixer 343 ofextractor E4 342. In addition extractant solution 302 is pumped via pump306 to mixer 343. After mixing with aqueous from outlet 336 of extractorE3, the mixture flows through the settler 345 of the extractor E4 342.The mixture is allowed to separate into an aqueous phase and anextractant phase. The extractant phase is withdrawn by an overflow weirinto a collector 344 having an outlet 345 and then flows via anextractant line into the mixer 332 of extractor E3 341. Aqueous leavesextractor E4 342 via outlet 346 to an oil water separator 347 and thento a sand filter 348 to obtain purified water 349.

Sulfate is stripped from the extractant in strippers 366, 371, 376, 381.Strippers 351 and 356 serve to remove metal ions with a charge of about+3 as metal sulfate product 355. Extractant, floc and entrained aqueousalso flow to the mixer 352 of stripper 351 from the outlet 315 of dualdecanter 314. Inflow of extractant to mixer 352 mixes with aqueousinflow from stripper 356 and the mixture in stripper 351 is separatedinto an aqueous MSO4 product 355 and an extractant that flows to a mixer357 in stripper 356. Note that all the strippers have internal recyclelines 303 for aqueous recycle flow. The incoming extractant fromstripper 351 is mixed in mixer 357 with sulfuric acid 353 (in thisexample about 50%) to make the mixture have a very low pH so that themetal is stripped from the extractant and driven into the aqueous phase.Extractant leaves stripper 356 and flows via an extractant line 305 tomixer 367 of sulfate stripper 366. In mixer 367 aqueous flow fromstripper 371 mixes with the incoming extractant from stripper 356.Sodium sulfate 368 that is a very useful product flows from the outletof stripper 366. An extractant line 305 from stripper 366 provides forflow of extractant to the mixer 372 of stripper 371. Aqueous fromstripper 376 flows to the mixer 372 and is mixed with the incomingextractant from stripper 366. Extractant further stripped of sulfatesexits stripper 371 and flows to the mixer 377 of stripper 376, here itis mixed with incoming aqueous from stripper 381. Extractant furtherstripped of sulfates exits stripper 376 and flows to the mixer 382 ofstripper 381, here it is mixed with incoming carbonate 383 and the like.The extractant essentially stripped of metals and sulfates flows fromstripper 381 to an extraction solution storage tank 303.

Referring now to FIG. 4, the present invention provides some designedfeatures that are specific to assist flowing and maintaining the flow ofthe metal oxide/hydroxide colloids, flocs and particulate slurries. Thedetailed schematic of FIG. 4 shows certain elements of the designedflotation device to assist in maintaining the flow of extractant phaseover the settler overflow weir while not interfering with aqueous phaseclarification. The flotation device flow pattern is also illustrated tovisualize the movement of the metal ion colloids, flocs and/orparticulate slurry.

The apparatus shown in FIG. 4 represents a typical mixer-settlerconfiguration. The apparatus has a mixer compartment 404 that has arotating mixer impeller 402 which creates a suction that pulls the AMDwater and extractant phase 406 into the mixer compartment 404 where itmixes the AMD water and extractant phase 406 thoroughly. The mixingforces the newly formed aqueous-extractant phase emulsion 424 over thetop of the mixer compartment 404 and under the underflow weir 409 intothe settler compartment 420. Here the mixed phase emulsion 424 isallowed sufficient time to break and split into 2 separate phases. Theupper layer of the settler compartment 420 contains the extractant phasewith floc and ions 422. The lower layer contains the aqueous phase 426that has been treated and depleted of ions. The bottom of the settlercompartment 420 contains an outlet 438 for the aqueous phase.

At the end of the settler compartment 420, there is a floc weir 430 thatfacilitates the extractant-floc phase movement. This weir contains afloc or slurry entrance surface positioned at typically about 30° fromthe vertical angle. The smooth entrance ramp 432 and exit ramp 434provide the least resistance for the extractant phase-floc phase to moveover the surface of the floc weir 430. And what is more, this angle offloc approach and exit, combined with the length of the ramps to use theinherent internal colligative property of the floc, which is gel-like inconsistency, to enable the exiting floc to literally pull the enteringfloc over the smooth weir. In this manner, floc flow is maintainedcontinuous or semi-continuous. Literally, the weight of the falling flocfilm and associated extractant phase pulls the entering extractant-flocphase layer to and over the weir. The stronger the internal colligativestrength the shorter the ramp can be for a given fluid viscosity. Bothan entrance angle and an exit angle are important design parameters asshown in the figure. Typically both of the angles can be the same ordifferent and can be from about 10° to about 80°, more preferably fromabout 15° to about 45°. At the right hand side of the figure, therounded lip 436 of the floc weir 430 enables efficient flow over theweir while the rounded bottom 435 of the floc weir 430 enables reliablegravity-driven break-off of the exiting floc containing extractant phaseliquid film that creates a flow pattern that allows the overflowingextractant phase to be dripped into the extractant solution where itexits the device via the extractant trough or launderer to the nextstage. Since the controls are typically performed via gravitationalforce, the drip point 437 is at the lowest point of the floc weir exitslide and the metal ion colloid thickened phase will not linger orattach to the sides of the weir or the tank.

An optional standpipe 450 that may be adjustable in height may be usedto control the outflow of aqueous. The standpipe 450 may be internal tothe settler compartment 420 as shown in FIG. 4 may be external as isknown in the art (not shown).

Referring now to FIG. 5, a further component of the invention is aextractant phase floc flow guide for the settler and/or decanter. Theupper right hand corner of the drawing shows a 3-dimensional view of ablock shaped flow guide that is also adjustable, it may be made ofdifferent layers of flat stock fixed on top of each other for thicknesstunability of the block to allow optional flow control for differenttypes and thicknesses of floc or may be a solid block. The device ishereinafter called a flow guide. The flow guide may be in the form of ablock shaped flow guide 510 a or as a plate shaped flow guide 510 b. Forfurther discussion below, the designation flow guides 512 a, 512 b, 512c will be used.

FIG. 5 also shows an expanded version of the extractor decanter systemfrom FIG. 3 but the larger scale showing more detail. The view is fromthe top and shows the flow guides installed into the “T” designedmixer-settler described above (FIG. 3). When acid mine drainage waterenters inlet 502 of mixer 514 in extractor 508 it is mixed withextractant solution and flows over or through a weir 511 to a settler516. The flow guide 512 a aids in maintaining the flow of extractantphase-floc (typically metal ion, hydroxide, sulfide containing) as itthickens during its travel along the settler 516 and over a floc weir517 where the floc exits the settler 516 and flows into a decanter 530.Aqueous and extractant entering the inlet of mixer 524 of extractor 520is mixed and flows over or through a weir 513 into a settler 526. Theflow guide 512 b aids in maintaining the flow of extractant phase-floc(typically metal ion, hydroxide, sulfide containing) as it thickensduring its travel along the settler 526 and over a floc weir 527 wherethe floc exits the settler 526 and flows into a decanter 530. Theinflowing floc from floc weir 517 and floc weir 527 combine in decanter530 further separating into an extractant phase containing floc andaqueous phase. The flow guide 512 c aids in guiding and maintaining theflow of the relatively thick floc over the floc weir 537 to outlet 540.The aqueous outlet pipes 519 a, 519 b, and 519 c provide individualcontrol of aqueous phase level and floc thickness in settler 516,settler 526 and decanter 530 typically by adjustment of the height ofthe respective pipe. The flow guides 512 a, 513 b, 513 c are placed tohave an underflow channel not visible in this top view, see FIG. 5A fora side view. Thus the aqueous phase flows under the flow guides thatsimultaneously block the flow of extractant phase and floc to theaqueous outlet pipes 519 a, 519 b, 519 c. The arrows show typical liquidflow directions.

FIG. 5 shows the geometry of the flow guides and where the flow guide islocated in the settler or decanter. The flow guide is designed to narrowthe settler liquid flow channel most preferably only near the topsurface of the settler so that it only increases the velocity in thefloc-loaded extractant phase and thus help in moving and for thickeningthe metal oxide colloids that are the floc and are located above thewater phase. The flow guide combines with the flotation overflow weirdevice described above, to dramatically increase the flow of metalcolloids floc particulates extractant phase into the next mixingchamber. The flow guide, which can also be constructed as just as aninternal wall or partition that is raised at the bottom to allow theaqueous phase to under-flow it to allow continuous removal of aqueouslayer via a suitable weir sump or stand pipe. Alternatively the flowguide can be substituted with a thin plate of material as shown.

FIG. 5A shows a side view of the mixer 524 and settler 526 of FIG. 5. Itshows the geometry and location of the flow guide 512 b in settler 526.The extractant phase and AMD water 522 a are pulled into the mixingcompartment 524 by the suction of the rotating mixer impeller 524. Inthe mixer compartment 524, the AMD water and extractant phase 522 a mixthoroughly. The mixing forces the newly formed aqueous-extractant phaseemulsion 527 a over the top of the mixer compartment 524 (in someembodiments the flow is under a weir 513) into the settler 526. In thesettler 526, there is a flow guide 512 b which moves the extractantphase and floc 525 along to the next mixing chamber. The figure showsthe flow guide 512 b constructed as an internal wall partition that israised at the bottom to allow space 512 b 3 for the aqueous phase 522 cto underflow the flow guide 512 b to allow continuous removal of aqueousphase 522 c via a stand pipe 519. Floc 522 b flows out over the flocweir 527

FIG. 5B shows an alternative version of FIG. 5A using a block 510 a as aflow guide. In this embodiment, the flow guide 512 b consists of anupper channel narrowing block located in the settler 526. The blockshaped flow guide 512 b is adjustable and it may be made of differentlayers of flat stock fixed on top of each other for thickness tunabilityof the block 510 a to allow optional flow control for different typesand thicknesses of floc or may be a solid block. In this figure, thereis an opening 519 b at the bottom of settler 526 which allows for theremoval of aqueous phase 519 c. Floc 522 b flows over the floc weir 527.

FIG. 5C shows the flow configuration of the E1 mixer compartment of FIG.17. The extractant phase and floc 541 from E2 decanter flow over thefloc weir 540 into the extractant trough or launderer 542. Theextractant phase and floc 541 are pulled into the mixing compartment 544along with the incoming AMD water 543 by the suction of the rotatingmixer impeller 545. In the mixer compartment 544, the AMD water 543 andextractant phase 541 mix thoroughly. The mixing forces the newly formedaqueous-extractant phase emulsion 547 over the top of the mixercompartment 544 and under the underflow weir into the E1 settlercompartment 546.

Gravity Benefits

Depending on the geographical location of the process unit, the watertreatment process according to the invention could be arranged in seriesto take advantage of the slope of a hill and make the gravitational flowmore efficient by providing steeper descents and deeper extractant phaseliquid thickener at over-flow weirs without the need to construct suchgradients.

Stage-to-Stage Loaded Extractant Phase Transfer Designs

When the unit extraction contactors configurations are in series, thereis an advantage to introduce the extractant solution without passing thefloc through piping to the mixer from the bottom. Though bothfloc-loaded modes work, most preferred is to avoid sending the flocthrough smaller I.D. piping to accomplish stage to stage transfers. Thisfeature avoids having the thickened extractant phase slurry or floc flowthrough piping and thereby run the risk of pipe or drain pluggage.However, when the extract material is introduced into a mixer (e.g. S1-Mfrom E1D) from the top, prevention of extractant flowing over into thestripper settler without proper mixing is critical to prevent shortcircuiting of the extractant phase flow. Sufficient mixing time of thefloc in the strip mixer is required to allow enough time and access toaqueous strip acid to dissolve the colloid/floc back into a conventionalmetal ion sulfate solution. Only enough time is needed to break theflocs affiliation with E-phase as continued hydrolysis reactions cancontinue in the internally recycled aqueous phase. This short circuitingis prevented by the trough or hopper feeding of the floc feed into themanifold beneath the mixer which feeds flocks, slurries, and fluids tothe mixing compartment by suction from the impeller. These designs thatare discussed below are readily implemented and do not require complexparts.

Referring now to FIG. 6, the most preferred design for stage-to-stagefloc transfer is to utilize an underflow weir 409 similar to the onethat is separating the mixer compartment 404 and settler compartment420. FIG. 6 shows the underflow inlet weir 408 position for introducingextractant floc solution 422 from the top of the mixer compartment 404and without piping. The rotating mixer impeller 402, while in action,creates a suction that pulls freshly added floc-loaded E-phase 422 tothe mixer 402, where it mixes the extractant solution 422 and inletaqueous phase 426 thoroughly while the underflow weir 409 preventpremature exit of the extractant 422 across the top of the mixercompartment 404 without first being thoroughly mixed with the inletaqueous phase 426.

Referring now to FIG. 7, another design, according to another aspect ofthe invention, is to place an extractant solution introduction inlet 480located inside the mixer compartment 404. The inlet 480 achieves asimilar effect as the underflow weir 409 by insuring thorough extractantsolution 422 and aqueous phase 426 mixing and reactor residence time. Inaddition, the bend elbow design on the bottom of the extractant phaseinlet line 480 (piping or tube) prevents feed water 426 back-flowinginto the tube which can avoid problems due to solids formation andpotential pluggage. The rotating mixer impeller 402 is most preferablydesigned to create a suction that brings the extractant solution 422 tothe bottom of the mixing compartment 404 and thereby promotes thoroughmixing with feed water 426.

Referring now to FIG. 8, another preferred design useful to enableprocessing of extractant phase flocs and slurries when introducingextractant solution with minimum piping is to include a storage chamber442 before the mixing compartment 404 and to introduce the extractantsolution 422 from the bottom of the mixing compartment 404 instead offrom the top. The inverted “y” shaped design feature at the bottom ofthe storage chamber 442 minimized aqueous phase 426 back-flowing intothe extractant storage chamber 442 and thereby avoids slurry pluggage ofthe extractant phase inlet line. This design also provides some inlinecontact between extractant 422 phase and aqueous phase 426.

FIGS. 9, 10, 11, and 12 provide typical extraction and strippingMcCabe-Thiele diagrams that were measured experimentally for candidateextractant formulations. These plots enable the user to determine thebest mode of operation of the invention with respect to the optimal E/Aratios and the number of stages needed for both the extraction andstripping of sulfate ions. These plots also identify the concentrationof the sulfate product produced and the level of residual sulfateremaining in the purified water at a selected set of operatingconditions.

Referring now to FIG. 13, the embodiment of the sulfate circuitcomprises four extraction stages for metal cation and sulfate anionco-extraction and four stripping stages for sulfate stripping andextractant regeneration. The overall process flow diagram is shown inFIG. 13.

The apparatus has a Na₂SO₄ solution (2000 ppm) inlet 1300 and variousinterconnections for aqueous flow referred to as internal recycle lines1381, aqueous lines 1383, and extractant lines 1385. Sulfate is strippedfrom the extractant in strippers S1-SO4 1339, S2-SO4 1338, S3-SO4 1337,and S4-SO4 1336. In addition, the extractant phase is regenerated in thestrippers. Note that all the strippers have internal recycle lines 1381for aqueous recycle flow. Aqueous phase is depleted of ions in E1decanter 1340 and extractors E1 1341, E2 1341, E3 1343, and E4 1343.

The Na₂SO₄ solution flows via pump 1315 and line 1302 to the mixingcompartment of the first extractor E1 1341. In addition, extractantphase and emulsion from extractor E2 1342 flows to extractor E1 1341 vialine 1306. The extractant phase and emulsion exits extractor E1 1341 vialine 1303 and enters E1 decanter 1340. In the decanter, the emulsion isallowed enough time to break and separate into an aqueous and extractantphase. The extractant phase flows to S1-SO4 1339 via line 1305 while theaqueous is withdrawn and sent to extractor E1 1341 via line 1322. Theextractant phase flows into S1-SO4 1339, exits via line 1307, and entersS2-SO4 1338. The aqueous phase exits S1-SO4 1339 as the Na₂SO₄concentrate product 1350 via line 1328. The extractant phase flows intoS2-SO4 1338, exits via line 1313, and enters S3-SO4 1337. The aqueousphase exits S2-SO4 via line 1309 and flows into S1-SO4 1339. Theextractant phase flows into S3-SO4 1337, exits via line 1315, and entersS4-SO4 1336 along with the incoming carbonate solution 1340 which flowsvia pump 1335 and line 1319. The aqueous phase exits S3-SO4 via line1311 and flows into S2-SO4 1338. The extractant phase flows throughS4-SO4 1336, exits via line 1321, and returns to the regeneratedextractant solution tank 1360. The aqueous phase leaving S4-SO4 1336exits via line 1317 and flows into S3-SO4 1337.

The aqueous phase withdrawn from E1 decanter 1340 flows into extractorE1 1341 via line 1322. The extractant phase leaving E1 decanter 1340exits via line 1305 and flows into S1-SO4 1339. The aqueous phase flowsthrough E1 1341, exits via line 1304, and flows into extractor E2 1342.The extractant phase exits E1 1341 via line 1303 and flow into E1decanter 1340. The aqueous phase flows through E2 1342, exits via line1310, and flows into extractor E3 1343. The extractant phase leaving E2exits via line 1306 and flows into E1 1341. The aqueous phase flowsthrough E3 1343, exits via line 1312, and flows into extractor E4 1344.The extractant phase exits E3 1343 via line 1310 and flows into E2 1342.The aqueous phase flows through E4 1344, exits via line 1318, and flowsinto the O/W separator 1380. The extractant phase exits E4 1344 via line1314 and flows into E3 1343. The purified water 1370 exits the O/Wseparator 1380 via line 1320. The extractant phase in the O/W separator1380 exits via line 1324 and flows into E4 1344 along with aqueous phasefrom E3 1343 which flows via line 1312 and additional extractant phase1360 which flows via pump 1325 and line 1316.

Referring now to FIG. 14, this figure illustrates acid mine drainagewater purification process flow Scheme 1. Process flow Scheme 1 providesthe ability to separate the metal ion components of the acid mine watersinto separate products as desired. FIG. 14 illustrates how this isaccomplished for two groups of metal ions, M and N. Given thisinformation, one skilled in the art would be readily able to proformmore that two separations by adding additional extraction or metal ionstrip circuits. This capability is described as follows using FIG. 14which deploys process flow scheme 1. More than two extraction stagescould be used for each extraction or strip circuit, but two were foundsufficiently effective in the current invention for most needs, howeverthree or four stages are preferred for each extraction or strip circuit.

FIG. 14 illustrates Scheme 1 and comprises two extraction stages for Mmetal cation and sulfate anion co-extraction, two extraction stageslinked in series to the first two stages for N metal cation and furthersulfate anion co-extraction, two (preferably three) acid strippingstages for M-metal stripping and two (preferably three) acid strippingstages for N-metal stripping, and four (preferably 4 to 6) strippingstages for sulfate stripping with simultaneous extractant regeneration.

The apparatus has an AMD feed water inlet 1400 and variousinterconnections for aqueous flow referred to as internal recycle lines1483, aqueous lines 1481, and extractant lines 1482. Sulfate is strippedfrom the extractant in strippers S1-SO4 1413, S2-SO4 1412, S3-SO4 1411and S4-SO4 1410. In addition, the extractant phase is regenerated in thestrippers. Note that the metal and sulfate strippers have internalrecycle lines 1483 for aqueous recycle flow. Aqueous phase is depletedof ions in E1-MSO4 Decanter (D) 1416, E1-NSO4 Decanter (D) 1424, andextractors E1-MSO4 1418, E2-MSO4 1419, E1-NSO4 1420, and E2-NSO4 1422.

The AMD feed water 1400 flows via pump 1475 and line 1403 into E1-MSO4mixer 1418 along with extractant phase and floc from E2-MSO4 via line1425. The aqueous phase (raffinate) exits E1-MSO4 1418 via line 1444 andflows into E2-MSO4 1419. The extractant phase exits E1-MSO4 1418 vialine 1426 and flows into E1-MSO4 Decanter (D) 1416. The aqueous phaseleaving E1-MSO4 Decanter (D) 1416 flows into the raffinate return 1485via line 1441. Raffinate Return from any Decanter of the inventionrepresents a relatively small amount of aqueous phase flow thatcorresponds to that volume of physically entrained aqueous phase thataccompanied the extractant phase flow as the latter exited its extractor(in this case E1-MSO4). It is an important and unique feature of thefloc based liquid liquid extraction technology that such physicallyentrained aqueous flows exist. These small aqueous flows exist due tothe requirement that the floc not be allowed to thicken excessively inthe extractor to avoid it thickening too much there and eventuallysolidify and plug up the setter and thereby becoming retained by theextractor. This problem is avoided by the technology by operating suchto keep the extractant phase layer thin in each extractor so as toreduce the residence time of the extract layer short. The thin nature ofthe extractant phase then allows some co-flow of aqueous layer as themixture approaches the floc over flow weir. This aqueous flow is easilycollected in the decanter and recycled as shown in the figures.

The extractant phase leaving E1-MSO4 Decanter (D) 1416 flows intoS1-MSO4 mixer 1415 via line 1427. The aqueous phase exits S1-MSO4 1415as the M-metal sulfate product (MSO4) 1491 via line 1440. The extractantphase exits S1-MSO4 1415 via line 1428 and flows into S2-MSO4 mixer 1414along with the 50 wt % sulfuric acid solution 1490, (or other strip acidfeed which is 2-70% in concentration, and preferably 15 to 50%concentration), which flows via pump 1474 and line 1405. The aqueousphase leaving S2-MSO4 1414 flows into S1-MSO4 1415 via line 1439. Theextractant phase leaving S2-MSO4 1414 flows into S1-SO4 1413 via line1429 along with aqueous phase from S2-SO4 1412 which flows via line 1436and extractant phase from S2-NSO4 1422 which flows via line 1456. Theaqueous phase exits S1-SO4 1413 as the Na₂SO₄ product 1489 via line1437. The extractant phase exits S1-SO4 1413 via line 1430 and flowsinto S2-SO4 mixer 1412 along with aqueous phase from S3-SO4 via line1435. Carbon dioxide (CO₂) 1486 exits S1-SO4 1413 via line 1438, whichis optionally captured and used as a co-product or vented. The aqueousphase leaving S2-SO4 1412 flows into S1-SO4 mixer 1413 via line 1436.The extractant phase leaving S2-SO4 1412 flows into S3-SO4 mixer 1411via line 1431 along with aqueous phase from S4-SO4 1410 which flows vialine 1434. The aqueous phase exits S3-SO4 1411 and flows into S2-SO4mixer 1412 via line 1435. The extractant phase exits S3-SO4 1411 andflows into S4-SO4 mixer 1410 via line 1432 along with 1 to 25% Na₂CO₃solution, preferably the 10 to 20% Na₂CO₃, and most preferably 13-17%Na₂CO₃ solution 1488 which flows via pump 1473 and line 1406. Theaqueous phase leaving S4-SO4 1410 flows into S3-SO4 mixer 1411 via line1434. The extractant phase leaving S4-SO4 1410 returns to the extractantsolution storage tank 1487 via line 1433.

The aqueous phase leaving E1-MSO4 1418 flows into E2-MSO4 mixer 1419 vialine 1444 along with extractant phase from the extractant storage tank1487 which flows via pump 1472 and line 1401. The extractant phaseleaving E1-MSO4 1418 flows into E1-MSO4 Decanter (D) 1416 via line 1426.The aqueous phase exits E2-MSO4 1419 and, for the case of using only two“M” extraction stages, flows into E1-NSO4 stage 1420 via line 1445 alongwith extractant phase from E2-NSO4 1421 which flows via line 1448. Theextractant phase exits E2-MSO4 1419 and flows into E1-MSO4 1418 via line1425. The aqueous phase leaving E1-NSO4 1420 flows into E2-NSO4 mixer1421 via line 1446. The extractant phase leaving E1-NSO4 1420 flows intoDecanter (D) 1424 via line 1453. The aqueous phase exits E2-NSO4 1421via line 1447 and flows into O/W separator 1417. The O/W separator 1417effluent exits via line 1443 as the purified water product 1493.Depending on the requirements for use or environmental release of thiswater it can be released, deodorized, and/or filtered for solidparticulate removal. The recovered low flow of extractant phase exitsthe O/W separator 1417 via line 1442 and flows into the extractantsolution recycle tank 1492. The aqueous phase leaving Decanter (D) 1424flows to the raffinate return 1484 (see above for definition ofRaffinate Return) via line 1454. The extractant phase leaving Decanter(D) 1424 flows into S1-NSO4 1423 via line 1452 along with aqueous phasefrom S2-NSO4 1422 which flows via line 1449. The aqueous phase exitsS1-NSO4 1423 via line 1451 as the N-metal sulfate product 1495. Theextractant phase exits S1-NSO4 1423 via line 1450 and flows into theS2-NSO4 mixer 1422. The stripped extractant phase from S2-NSO4 1422settler and flows to the S1-SO4 sulfate stripper 1413 via line 1456.

Referring now to FIG. 15, this figure illustrates acid mine drainageprocess flow scheme 2. Process flow scheme 2 comprises four extractionstages for metal cation and sulfate anion co-extraction, two acidstripping stages for metal stripping, and four stripping stages forsulfate stripping and extractant regeneration. The overall process flowdiagram is shown in FIG. 15.

The apparatus has an AMD feed water inlet 1500 and variousinterconnections for aqueous flow referred to as internal recycle lines1591, aqueous lines 1592, and extractant lines 1593. Sulfate is strippedfrom the extractant in strippers S1-SO4 1539, S2-SO4 1538, S3-SO4 1537,and S4-SO4 1536. In addition, the extractant phase is regenerated in thestrippers. Note that the metal and sulfate strippers have internalrecycle lines 1591 for aqueous recycle flow. Aqueous phase is depletedof ions in E1 Decanter (D) 1542 and extractors E1 1543, E2 1544, E2Decanter (D) 1545, E3 1546, and E4 1547.

The AMD feed water 1500 flows via pump 1555 and line 1530 into E1 mixer1543 along with extractant phase and floc from E2 1544 via line 1525 andaqueous phase from E1 D 1542 via line 1524. The aqueous phase exits E11543 via line 1526 and flows into E2 mixer 1544. The extractant phaseexits E1 1543 and flows into E1 D mixer 1542 via line 1507. The aqueousphase leaving E1 D 1542 flows back to E1 mixer 1543 via line 1524. Theextractant phase leaving E1 D 1542 flows to S1-MSO4 mixer 1541 via line1508 to line 1535 along with extractant phase from E2D 1545 via line1506 and extractant phase from O/W separator 1598. The aqueous phaseexits S1-MSO4 1541 as the metal sulfate concentrate product 1596 vialine 1548. The extractant phase exits S1-MSO4 1541 via line 1509 andflows into S2-MSO4 mixer 1540 along with the 50 wt % sulfuric acidsolution which flows via pump 1585 and line 1532. The aqueous phaseleaving S2-MSO4 1540 flows to S1-MSO4 mixer 1541 via line 1523. Theextractant phase leaving S2-MSO4 1540 flows to S1-SO4 mixer 1539 vialine 1510. The aqueous phase exits S1-SO4 1539 as the Na₂SO₄ product1592 via line 1533. The extractant phase exits S1-SO4 1539 and flowsinto S2-SO4 mixer 1538 via line 1511. The aqueous phase leaving S2-SO41538 flows into S1-SO4 mixer 1539 via line 1522. The extractant phaseleaving S2-SO4 1538 flows into S3-SO4 mixer 1537 via line 1512. Theaqueous phase exits S3-SO4 1537 and flows into S2-SO4 mixer 1538 vialine 1521. The extractant phase exits S3-SO4 1537 and flows into S4-SO4mixer 1536 via line 1513 along with the 15 wt % Na₂CO₃ solution whichflows via pump 1595 and line 1534. The aqueous phase leaving S4-SO4 1536flows to S3-SO4 mixer 1537 via line 1520. The extractant phase leavingS4-SO4 1536 is regenerated and returns to the extractant solutionstorage tank 1597 via line 1514.

The aqueous phase withdrawn from E1 D 1542 via line 1524 flows into theE1 mixer 1543. The aqueous phase exits E1 1543 via line 1526 and flowsinto E2 mixer 1544 along with the extractant solution from the storagetank 1597 which flows via pump 1575 and line 1531. The extractant phaseexits E1 via line 1507 and flows into E1 D mixer 1542. The aqueous phaseleaving E2 1544 flows into E3 mixer 1546 via line 1527 along withaqueous phase from E2D 1545 via line 1528 and extractant phase from E41547 via line 1504. The extractant phase leaving E2 1544 flows into E1mixer 1543 via line 1525. The aqueous phase exits E3 1546 via line 1529and flows into the E4 mixer 1547 along with extractant phase from thestorage tank 1597 which flows via pump 1565 and line 1501. Theextractant phase exits E3 1546 via line 1505 and flows into E2D mixer1545. The aqueous phase leaving E4 1547 flows to the O/W separator 1598via line 1502. The extractant phase leaving E4 1547 flows to E3 1546 vialine 1504. The O/W effluent exits via line 1503 as the purified water1599. The extractant phase in the O/W separator 1598 exits via line 1535and flows into S1-MSO4 mixer 1541.

Referring now to FIG. 17, this figure illustrates a presently preferredconfiguration: M are mixers, E1 to E4 are extractor units withassociated decanters and weirs having flow control guides, and S1-SO4 toS4-SO4 are the sulfate strippers. Typically, the extractor floc weirshave entrance and exit ramps with a rounded lip and bottom. Theconfiguration shows a water purification circuit 1801 m, a metal stripcircuit 1803, and a sulfate strip circuit and extractant phaseregeneration 1805. Metal recovery is in metal stripper S1M where a metalsulfate concentrate product is obtained. Note that the metal and sulfatestrippers have internal recycle lines for aqueous recycle flow.

The apparatus has an AMD feed water inlet 1800 and variousinterconnections for flow referred to as aqueous lines 1853 andextractant lines 1855. The AMD feed water 1800 flows via line 1802 intoE1 mixer (M) 1832 along with extractant phase and floc from E2 decanter1823. The mixture flows past flow guide 1834 and over weir 1835 into E1decanter 1833. In E1 decanter 1833, the floc thickens and maintainscontinuous flow due to flow guide 1836 and floc weir 1835. The aqueousphase flows under flow guides 1834 and 1836 and returns to E2 M 1822 viatwo standpipes and line 1858 in E1 decanter 1833. The extractant phaseand floc enters S1-M mixer (M) 1840 along with the acidic aqueous phasevia line 1856 from S2-M 1841. The floc is disintegrated by the acidicaqueous phase and forms acidic aqueous and extractant phases. Theaqueous phase leaving S1-M 1840 is the metal sulfate concentrate product1811 and can be collected via line 1808 in a storage vessel. Theextractant phase leaving S1-M 1840 flows into S2-M mixer (M) 1841 vialine 1818 along with the sulfuric acid solution 1813 which flows vialine 1810. The acidic aqueous phase exits S2-M 1841 via line 1856 andflows into S1-M mixer (M) 1840. The extractant phase leaving S2-M 1841flows into S1-SO4 mixer (M) 1842 with aqueous phase from S2-SO4 1843.The aqueous phase exits S1-SO4 1842 as the sulfate concentrate product1815 and is collected via line 1812. The extractant phase leaving S1-SO41842 flows into S2-SO4 mixer (M) 1843 with aqueous phase from S3-SO41844. The aqueous phase exits S2-SO4 1843 and flows into S1-SO4 mixer(M) 1842. The extractant phase exits S2-SO4 1843 and flows into S3-SO4mixer (M) 1844. The aqueous phase leaving S3-SO4 1844 flows into S2-SO4mixer (M) 1843. The extractant phase leaving S3-SO4 1844 flows intoS4-SO4 mixer (M) 1845 along with the carbonate solution 1817 which flowsvia line 1814. The aqueous phase leaving S4-SO4 1845 flows into S3-SO4mixer (M) 1844. The extractant phase exits S4-SO4 via line 1816 as theregenerated extractant solution 1819 and is collected in a storagevessel.

The aqueous phase withdrawn from E1 decanter 1833 via line 1858 and twostandpipes flows into extractor E2 mixer (M) 1822. The extractant phasein E1 decanter 1833 flows over flow weir 1835 and flows into S1-M 1840.The aqueous phase flows through E2 1822 and exits via flow guides 1824and 1857 and two standpipes to line 1859. The extractant phase and flocleaving E2D 1823 flows into E1 mixer (M) 1832 along with AMD feed water1800 which flows via line 1802. The aqueous phase flows into E3 mixer(M) 1821 via line 1859 with extractant phase from E4 1820. The aqueousleaving E3 1821 flows into E4 mixer (M) 1820 via line 1860 with freshextractant phase 1807 which flows via line 1804. The extractant phaseleaving E3 flows into extractor E2 mixer (M) 1822 via line 1861. Theaqueous phase flows through E4 1820 and exits via line 1806 as thepurified water product 1809. The extractant phase leaving E4 1820 flowsinto E3 mixer (M) 1821 via line 1862.

To demonstrate that the selected process of the invention was scaleable,a field demonstrating unit was constructed with process flow diagramsimilar to that of FIG. 15 but sized to be operable at 5 to 40 gal/min,and optimal at 12 to 30 gal/min acid mine drainage water feed rate. Thefield unit was the size of two tractor trailer beds and located on anabandoned mine site at St. Michael, Pa. where a continuous acid minedischarge water is flowing at a rate of 10,000 gallons/min so thatsufficient feed water was available for 24/7 operation. These trialslasted for 3 months where the water purification process chemistrypreviously described was confirmed at this larger scale. The E/A ratiowas varied in this three months of testing over the range of 1/2 to 1/10and where the preferred ratio was 1/3 to 1/8, and most preferred ratiowas the same as was found for the laboratory and pilot testing, or 1/4to 1/6. This testing demonstrated the robustness of the technology, itsgood economics, ease of operation, and water purification capability.

EXAMPLES

The following examples illustrate various aspects of the invention andare not intended to limit the scope of the invention in any way.

Example 1

The objective of this test was to determine the key parameters andranges necessary for a successful removal of contaminants from acid minedrainage water and accomplishing the separation of the recoveredcomponents while forming them into concentrates. The processconfiguration in FIG. 13 was constructed using clear chemical resistantPVC (CPVC) for the mixer-settler tanks, which had had an internal mixervolume of 186.7 cm³ and an internal settler volume of 401.5 cm³. Theflow rate used resulted in a 333.6 second residence time. Clear Tygontubing (0.25″ I.D.) was used for the piping. Cole-Parmer InstrumentCompany Master-flex L/S Peristaltic pumps and Dayton AC-DC series motormixers were used. The extractant solution formulation was prepared using15% Aliquat 134, 15% Exxon 10 (Isodecanol), and Calumet 400-500. Alltest conditions are given in the LLX Test Condition Key.

The process was started up, operated, and shut down in the followingmanner.

At start-up, the system was charged with aqueous solutions first, andeach mixer settler of the process was charged with approximately 50% ofits respective volume. The system is initially charged with 50% SulfuricAcid, followed by a 15% Sodium Carbonate solution. The more preferredmethod is to charge the strippers with 25% H₂SO₄ to stay away fromexcessive strong acid which can tend to third phase formation. Mostpreferred is to charge the system with actual M-SO₄ concentrate from aprevious run. Charging the system in this manner causes the extractantoverflow receiving compartments to partially fill. After this phase ofstart-up is complete the system is now ready for the extractantsolution.

The extractant solution must be fully acid stripped of Cl⁻ and thencarbonate loaded (1-25% Na₂CO₃), preferably 15% Na₂CO₃) before beingadded to the process. Although the introduction of extractant solutionis best achieved at the pilot and commercial scale levels using pumps,at the bench/lab scale this can be quickly achieved by manually pouringthe extractant solution into the mixer-settlers to the point of overflowinto the E-phase overflow compartments. After charging the flotationliquid-liquid extraction circuit with a sufficient volume of extractantsolution, there needs to be enough extractant solution left in the surgetank so that the process needs are met during normal operation. Thetotal volume of the extractant solution surge tank should be designedlarge enough so that it does not overflow during the operation of theflotation liquid-liquid extraction process and can be charged withextractant solution when the system is shut down between operations. Thesteady-state volume of the extractant solution in the surge tank is thenmonitored visually or electronically with level switches. This should bedone periodically so that the extractant solution surge tank volume canbe adjusted as needed to maintain sufficient extractant solution volumeto provide steady operation over extended periods, for example weeks,months and possible years. The stirrers for the mixers were then poweredup, adjusted and maintained at steady-state by the following procedures.

All of the mixers were set between 700-1700 rotations per minute (rpm).The mixers need at least 15 minutes to warm up, preferably 30 minutes.During this time the mixers are monitored and adjusted, usuallydecreasing the rotation rate in order to avoid excessive mixing.Excessive mixing is very undesirable. It can lead to problems such asspatter as well as the formation of fine emulsions that may be stable orthat require longer phase coalescence time in the settlers. Although anytype of stirring is sufficient enough to mix medium to low viscosityimmiscible fluids, disk or fin type stirrer pumps are preferred. Theyare both designed to pull the two fluids, aqueous and extractantsolution, into the mixing compartment from the upstream mixer settlers.The shearing blades of the mixers generate micro droplets that create avery high interfacial surface area that is critical to fast contaminantextraction and strip kinetics. Higher mixing speeds accommodate shorterresidence time of the fluid in the mixer and compensate forextractant/aqueous ratios other than 1:1. Although stirrer speeds thatresult in the mutual blending of only 20% of the two phases issufficient, a blend of at least 80% is preferred and if optimumconditions can be achieved 95-100%. Excessive mixing is suitable butless preferred if the resultant emulsion formed requires mixing for longperiods of time to disengage and break due to exceedingly fine dropletsize. Mixing conditions preferred by the invention is about 12-120seconds, preferably 30-90 seconds, and if optimum conditions can beachieved 45-seconds. The total hydraulic fluid residence time in themixer and the settler necessary for this process should be 10 times thatamount or approximately 15 minutes. Due to the low values and/orquantities of the contaminants present in the Acid Mine Drainage waterpurification process (Fe, Al, Se, Si, Mn, Zn, Ni, Co, Ca, SO₄ ⁻²), andthe very high flow rates of 10-10,000 gal/min (averaging about 500gal/min, but often variable) of water, conventional metals extraction byliquid-liquid extraction is not feasible because of the very largeequipment and E/A ratio requirements that would be required for thelarge aqueous flows involved in water purification. Although the mixingconditions can be either extractant phase continuous or aqueouscontinuous, the latter is the more typical since special startupconditions are not needed to achieve it.

After this initial loading of strip solutions and the warm up time forthe mixers are both complete, the system is now ready for the chargingof the acid mine drainage feed water and {the acid mine drainage feedpump was set to 72 ml/min at system start-up (test parameter goal foracid mine drainage feed stream is 100 ml/min)} the extractant phase feedstreams. The N-extractant (extraction circuit designed to extract+2metal ions) feed flow rate at system start-up is set to 18 ml/min. Theextractant surge tank for this process was a 4 L clear chemicalresistant PVC tank. Other feed agents are listed in the LLX TestCondition Key.

Operation, Control, Monitoring (approaching and maintainingsteady-state): The process is run for approximately 20 hours beforesteady-state is reached. This gives the extractant phase enough time tocycle through the process at least 3 times (assuming 6 L surge tank used@ 15 ml/min). The extractant phase contacts the feed first during themetal extraction stage and then gets metal stripped with Sulfuric Acid(50% H₂SO₄). After the Metal Strip Stage the extractant gets carbonate(15% Na₂CO₃) loaded to extract the sulfate ions and is now ready for useagain. The extractant phases flow scheme is also illustrated in FIG. 13.

Once the extractant has had enough time to cycle through the system atleast once approximately 80% of the extractant phase stays within thesystem and the remaining 20% stays in the surge tank. The acid minedrainage feed rate was increased in 10 ml/min increments until the testparameter of 100 ml/min was reached. Occasionally the system must be putin idle mode (park) until certain control issues can be adjusted, anychemical additions can be added, or any maintenance issue can beaddressed. When this is done all of the feed pumps are shut off and themixers are allowed to continue to circulate the process fluids. Forexample, for this test the process was parked for the following reasons:extractant feed levels running too low, recalibration of pumps needed,aqueous in extractant feed pump, etc. Once the process is said to havereached steady-state data readings can be collected and sampling can nowtake place.

The samples for this test were collected out of the metal and sulfatestrip extraction stages as well as the acid mine drainage Feed andRaffinate discharge lines(E1-E2-E3-E4-S1SO4-S2SO4-S3SO4-S4SO4-Feed-Raffinate). There were twosamples taken in 1 L bottles from the acid mine drainage drum per runfor analysis (sample with 2% Nitric Acid and “as is” sample).

Each extraction stage sample was taken from the aqueous phase of thefluid in the settler tank using plastic disposable Luer Lock syringes,and filtered using Serum Acrodisc 37 mm syringe filters with glass fiber(GF)/0.2 micron pores. This precautionary measure was taken to assure aminimal amount of organic phase within the sample. The samples weretaken within one hour of each other. There were four samples taken fromthis run. The data collection should be done in collaboration with thesample collection every hour, it consists of: pH, density, E:A ratio,chemical volume, extractant depth, and mixer tip speeds. Once the datahas been collected it will immediately be place inside of the labnotebook. The samples are then sent off for Inductive Coupled Plasma(ICP) and Inductive Coupled Plasma Mass Spectrometry (ICP-MS) analysis.

At the end of this test 37.2 gallons of acid mine drainage feed had beenused over the course of approximately 40 hours.

Example 1A (Stages Determination)

This example determined the effect of sulfate extraction efficiencyusing extractant phase formulation. The apparatus of FIG. 13 wasconstructed. The extractant phase was prepared by selecting componentsfrom Table 3, and certain components of these were selected and blendedin proportions given in Table 5. The test conditions used are given inTable 5.

The experiment used extractant solution from Table 5 to extract sulfatefrom 2000 ppm Na₂SO₄ solution. Since the extraction of sulfate anions ina Na₂SO₄ solution rely mainly on a concentration gradient, this set-upserved as “worst case” in sulfate anion removal of water purificationprocess. The extractant solution was treated with 15% Na₂CO₃ to load upthe extractant with carbonate anion, and then the extractant and Na₂SO₄solution were tested under different E/A ratio. The aqueous phase wasanalyzed using ion chromatography to determine the remaining sulfateconcentration. The ion chromatography results determined the amount ofsulfate removed at different E/A ratio. And those results were alsoplotted via McCabe-Thiele plot, which lead to the discovery of thestages needed for sufficient sulfate loading and stripping of the watertreatment process. The experiment also determined the range ofextractant needed for effective sulfate extraction, which preferred bythe invention is about 5-15%, most preferably 8-12%. The McCabe-Thieleresult is shown in FIG. 9 to FIG. 12.

To validate the results, a sulfate extraction process was set up withthe apparatus of FIG. 13. This process was designed to extract sulfateanion from a 2000 ppm Na₂SO₄ solution via a continuous circuit. Theresult from this process validated the proposed design and the samesulfate process apparatus was implemented into the overall process flowscheme. This also established a baseline test condition, which includesE/A ratio between 1/4-1/6 for sufficient sulfate extraction andstripping in three to four stages.

Example 2 (Run #2)

This example determined the effect of metal and sulfate extractionefficiency in a less contaminated stream using extractant phaseformulation. The apparatus of FIG. 14 was constructed. The extractantphase was prepared by selecting components from Table 3 blended inproportions giving in Table 5. The test conditions used are given inTable 5.

This experiment used extractant solution from Table 5 with processapparatus of FIG. 14 to perform water purification process with an acidmine drainage stream containing low contaminant levels. The process wasset up with the same number of extracting and stripping stages. Thecombination of high extraction efficiency of metal cation and sulfateanion and the low contaminant concentration in the acid mine drainagestream caused secondary emulsion at the third extraction stage. Theemulsion also stopped phase disengagement between extractant solutionand the raffinate. The extractant solution and raffinate emulsioncreated from the process behaved similarly to an emulsion of anextractant solution and deionized water. The emulsified raffinate fromthis process was left in the tank for 24 hours to prolong the settlingtime in an attempt to achieve better phase disengagement. However, therewas no visible phase separation after 24+ hours of settling time, andthe emulsion was not able to break until pH adjustment was made byaddition of sulfuric acid. It appears that mixing was excessive thusforming a stable microemulsion. Stable in that the disengagement fromthe emulsion was too slow. When mixer speed was controlled this problemdid not return.

Example 3 (Run #6)

This example determined the effect of acid mine drainage liquid-liquidextraction (LLX) system efficiency with shorter mixing residence timeusing extractant phase formulation. The apparatus of FIG. 14 wasconstructed. The extractant phase was prepared by selecting componentsfrom Table 3 blended in ranges given in Table 4. The test conditionsused are given in Table 5.

This experiment also used extractant solution from Table 5 with processapparatus of FIG. 14 to evaluate acid mine drainage water purificationprocess performance under shorter mixing residence time. The two-partexperiment tested different process conditions; the first part of theexperiment tested the treatment process with E/A ratio of 1/6 and mixingresidence time of 60 seconds, and the second part tested the treatmentprocess with E/A ratio of 1/8 and mixing residence time of 35 seconds.

The 60 seconds mixing residence time trial was behaving well on its own,and the consumption rates of Na₂CO₃ and H₂SO4 were determined in thisexperiment. The raffinate from individual stream was sent to ionchromatography and Inductively Coupled Plasma Mass Spectrometry (ICP/MS)to determine the final metal and sulfate concentration. These resultsand measurements helped to determine the final mass/molar balance of theoverall process. The 35 seconds mixing residence time trial also had asmooth operation. However, the raffinate's ion chromatography and ICP/MSresults from this trial turned out not fulfilling the performancerequirement. This allowed the determination of the mixing residence timeand E/A ratio limitation for the treatment system. The analyticalresults from these trials are shown in Table 6.

The result from the two-part experiment not only showed the limitationof E/A ratio and mixing residence time needed for sufficient processtreatment, but also determined the extractant solution input rate withrespect to the mixer size. If too much extractant solution pushed intostrippers, the mixer section would be overwhelmed and started pluggingup because of the excess extractant solution in the mixer. The excessextractant solution in the mixer had gel-like behavior and caused themixer to lose hydraulic suction. This would severely cripple the processbecause it was essentially equal to shutting down the process atmidpoint while more acid mine drainage water and extractant were beingpushed through the front end of the process. The discovery of thisparticular limitation allowed the determination of the maximumextractant solution flow rate into the strippers, thus provided theboundary condition needed for the process.

Example 4

This example determined the effect of water entrainment from extractantphase formulation, especially the contribution of chosen extractant anddiluent and their relationship. The test apparatus consisted of a seriesof graduated, capped vials used to perform batch evaluations of thedegree of water entrainment with respect to extractant phase formulationat extraction and strip conditions, and while preserving good yields ofsulfate ion extraction. The extractant phase was prepared by selectingcomponents from Table 3; blended in proportions given in Table 5. Thetest conditions used are given in Table 5.

This series of experiments used extractant solution from Table 5 todetermine the “best case” extractant phase formulation that yielded theleast water entrainment in the extractant phase. The experiments weredesigned by using Design Expert, version 6.0, and the goal for thisexperiment was to discover the extractant phase formulation range thatwould potentially produce the highest metal sulfate product withoutcompromising the sulfate extraction efficiency. There were fivedifferent concentrations of extractant solutions and each wasconditioned with 15% Na₂CO₃. The different extractant solutions werethoroughly mixed with acid mine drainage water with E/A ratio of 1/6.The mixtures were then left to settle and the settling time wasrecorded. A certain amount of extractant phase was transferred to acentrifuge to pull any entrained raffinate out from the extractantphase. The leftover raffinate was then transferred out and theextractant phase was then treated with 50% H₂SO₄ for metal extraction.Measurements such as phase disengagement time, amount of raffinateentrained, color of the extractant phase, etc. were taken during theexperiments and the results were entered into the Design Expertsoftware.

FIG. 16, from Design Expert version 6.0, showed schematically therelationship of the concentration of Aliquat and Isodecanol and theirimpact on post-phase separation water entrainment in the extractantsolution. The goal was to find a “best case” formulation and testcondition with respect to high metal sulfate production. This experimentnot only led to the discovery of the “best case” test condition neededfor producing high metal sulfate product, but also demonstrated theeffect of the rapidly declining metal sulfate concentration caused bywater entrainment. If the raffinate water was entrained in extractantsolution, even a small amount of water would reduce the metal sulfateconcentration dramatically and sacrificed the salability of the metalsulfate product. The diluted metal sulfate concentration would also needa long time to build back up to a desirable level.

Example 5 (Run #7, Scheme 3)

This example determined the effect of the optimum extractant phaseformulation found in the previous designed experiments (example 4) andits influence on metal sulfate product concentration. The apparatus ofFIG. 3 was constructed. The extractant phase was prepared by selectingcomponents from Table 3 and blended in proportions giving in Table 4.The test conditions used are given in Table 5.

This experiment used extractant solution from Table 5 with processapparatus of FIG. 3 to evaluate the extractant phase “best case”formulation with respect to amount of water entrainment in theextractant phase and the concentration increase in the metal sulfateproduct stream. The experiment was set up to test several different testconditions, and a set of measurements were conducted from each of thetest conditions to determine the metal sulfate product flow rate and theconsumption rate of sulfuric acid. These measurements were critical fordetermining the metal sulfate production rate and provide a morerealistic economic estimation for the field-scale process unit. Theresults from the experiment were then entered into the Design Expertsoftware, version 6.0 to determine the optimum test condition forproducing high concentration metal sulfate product stream. Testcondition range and the optimum test condition that would be needed toproduce high metal sulfate concentrations were determined. As determinedby the software, the E/A ratio would be increased from 1/6 to 1/4.35,with longer mixing residence time of 112.5 seconds. The result from thisexperiment will serve as the test condition for the next experiment, andthe metal sulfate product concentration would be compared with theprevious tests and a final extractant phase formulation could bedetermined.

Example 6 (Run #8)

This example determined the effect of the “best case” extractant phasetest condition found in the previous designed experiments (example 5)and its influence on metal sulfate product concentration. The apparatusof FIG. 3 was constructed. The extractant phase was prepared byselecting components from Table 5 and blended in proportions giving inTable 4. The test conditions used are given in Table 5.

With the previously determined test condition, the experiment ran over27 hours without major problems and was also able to provide valuableinformation on the metal sulfate product concentration. The flow rate ofmetal sulfate product was determined, and the ratio of metal sulfateproduct to acid mine drainage feed flow rate were about 1/5 or less.

The process also added sodium hydroxide solution into the last stage ofextraction in an attempt to extract out additional magnesium via pHcontrol. And some solid particles precipitated and settled in the bottomof the last extraction settler. This also indicated the potentialproduct that could be produced from this process and also the additionalimplementation to avoid solid blockage inside the transferring lines.The results for this experiment are shown in Table 8

This experiment not only showed the potential concentration of the metalsulfate product, but also the sensitivity of such product stream towardwater carry-over by the extractant solution. Even a small amount ofwater/raffinate carried over by the extractant solution would decreasethe concentration tremendously. And once the concentration dropped, itwould take a long time to build up the concentration back to thedesirable level. The diluted metal sulfate product stream would alsocreate faster harvesting flow rate, i.e. the ratio of metal sulfateproduct to acid mine drainage feed flow rate would increase to 1/3 ormore. The sudden surge of water/raffinate disrupted the level of theextractant and aqueous phase and caused domino effect to all thedownstream process. Therefore, it was essential to find a good controlmethod for metal sulfate stripper to guarantee the success of thisprocess.

Example 7 (Run #9)

This example determined the effect of the extractant phase testcondition with alternative diluent in the formulation and its influenceon metal sulfate product concentration. The apparatus of FIG. 3 wasconstructed. The extractant phase was prepared by selecting componentsfrom Table 5 and blending in proportions giving in Table 4. The testconditions used are given in Table 5, .see Run #9 in this table whereScheme 3 denotes the apparatus in FIG. 3.

This experiment evaluated′ the metal sulfate product concentration andthe presence of water/raffinate in the extractant solution withalternative diluent. Additional bench-scale testing showed thatincorporating aromatic diluent in extractant phase formulation wouldyield low water/raffinate entrainment in the extractant solution, thuslower the water/raffinate carry-over into the metal sulfate strippers.The extractant phase formulation with aromatic diluent behaved well inthe same process configuration. And the alternative extractant phaseformulation performed as expected and significantly decreased thewater/raffinate carry-over into the metal sulfate strippers.

The test condition of this experiment was kept at baseline processcondition, i.e. 60 seconds mixing residence time, E/A ratio=1/6. Testingwith the baseline process condition provided grounds for comparisonbetween different extractant phase formulation and the effect of usingdifferent diluent. The formulation with aromatic diluent showed themetal oxide colloids in extractant solution were flowing well and didnot create clumps of metal oxide floc. This definitely helped thetransfer efficiency of metal oxide colloids into metal sulfate strippersdue to the fluidity of the solution. Unfortunately, the process had somewater carry-over during part of the experiment, which decreased themetal sulfate concentration. However, once the process was back tosteady state condition, the metal sulfate concentration would build upquickly.

This experiment provided grounds for baseline comparison betweenextractant phase formulation effect with respect to aliphatic andaromatic diluent. This experiment had demonstrated the efficiency ofusing aromatic diluent for extractant phase formulation. However,aromatic diluent had long term material compatibility problem with thecurrent setup, which provided the basis for determining an optimumextractant phase formulation with proper diluent mix. The optimumextractant phase formulation would have low water/raffinate entrainmentin extractant phase while being compatible with easily accessedmaterials, such as PVC or fiberglass reinforced plastics.

Table 4. Extractant phase formulation table with typical minimum andmaximum extractant component. The extractant, modifier, and diluent canbe referred to

TABLE 3 Extractant phase Formulation Extractant Volume PercentageFormulation (%) Component Minimum Maximum Extractant 5 15 Modifier 2.315 Diluent 70 92.7

In a broader aspect of the invention the formulations may comprise thefollowing (in volume %):

Extractant: 0.5 minimum to 70 maximum;Modifier: 0 minimum to 95.5 maximum;Diluent: 0 minimum to 95.5 maximum.

The important criteria that is required within the narrow or broadermaterial limits is that a floc is formed that separates from the treatedaqueous phase and the loaded extractant phase.

Example 8

This example illustrates the determination of oil water separator andextractor setting time and the determination of the acid and baseconsumption rate during purification of acid mine drainage water usingthe invention.

The continuous process configuration in FIG. 9 was constructed usingclear, chemical resistant PVC for the mixer-settler tanks, which had aninternal mixer volume of 180 ml. This allowed the process fluids to havea residence time of 60 seconds with the ETA total (combined) flow rateof 180 ml/min. Additional equipment included, clear Tygon tubing for thepiping (0.375″ I.D.), Cole-Parmer Instrument Company Master-flex L/SPeristaltic pumps, and Dayton AC-DC series motor mixers. The extractantphase formulation was prepared using 15% (v/v) Aliquat 134®, 15% (v/v)Exxal 10° (Isodecanol), and 70% (v/v) Calumet 400-500 diluent. Thisdiluent is less than 1 wt % aromatics. Other test conditions are givenin Table 5.

Relative to example 4, the process alterations made were:

1. The M & N extraction box decanters were replaced with separatoryfunnels to allow fluid dimensions and sharp phase separator control.2. The M & N Sulfate product discharge lines fed directly into separatetanks.3. The M Sulfate decanter was modified by drilling angled holes into themixer overflow weirs to facilitate a discharge of equal or lesser heightof the E1 M Sulfate extractant phase flow.4. A peristaltic pump was installed to transfer the M, N metal ions(trivalent and divalent) and SO₄ loaded extractant phase(s) to theappropriate stripper.5. The extractant phase to aqueous acid mine drainage/extractant phaseratio was changed to 1:5.6. An oil/water separator was added to recover any extractant phase lostfrom the extraction operation to the aqueous raffinate exit stream ofthe E4 extraction stage.7. Double wide settler used for E3 and E4 (FIG. 3).8. To determine the impact of settling time during the extractionoperation.

The process was started up, operated, and shut down in the followingmanner:

Chemical Charging (Start-up): The system was charged with aqueoussolutions, with each mixer settler of the process charged withapproximately 50% of its respective volume. The strippers were initiallycharged with 5-50% (w/w) sulfuric acid. The extractant phaseregeneration mixer-settlers were charged with a 15% (w/w) sodiumcarbonate solution. By charging the system in this manner, typically theextractant overflow compartments fill to half full capacity. After thisphase of start-up is complete the system was now ready for theextractant phase.

Fresh extractant phase was optimally pre-cleaned prior to first chargingto the extraction system by fully acid stripped (0.1 50% v/v H₂SO₄,preferable 25%) and then carbonate loaded (15% w/w Na₂CO₃ (range0.5-30%) before being added to the process. Although the introduction ofextractant phase is best achieved at the pilot and commercial scalelevels using pumps, at the bench/lab scale this can be quickly achievedby pouring the extractant phase into the mixer-settlers to the point ofoverflow into the settler compartments. After charging the liquid-liquidextraction (LLX) circuit with a sufficient volume of extractant phase,there should be enough extractant phase left in the surge tank so thatthe process needs are met during normal operation. The total volume ofthe extractant phase surge tank should be large enough so that it doesnot overflow during the operation of the LLX process and can be chargedwith extractant phase when the system is shut down between operations.The steady-state volume of the extractant phase in the surge tank isthen monitored visually or electronically with level switches. Thisshould be done periodically so that the extractant phase surge tankvolume can be adjusted as needed to maintain sufficient extractant phasevolume to provide steady operation over extended periods, for exampleweeks, months and possibly years. The stirrers for the mixers were thenpowered up, adjusted and maintained at steady-state by the followingprocedures.

All of the mixers were set between 700-1700 rotations per minute (rpm).The mixers needed at least 15 minutes to warm up, preferably 30 minutes.During this time the mixers were monitored and adjusted, usuallydecreasing the rotation rate in order to avoid excessive mixing.Excessive mixing is very undesirable; it can lead to problems such asspatter as well as the formation of fine emulsions that require longerphase coalescence time in the settlers. Although any type of stirring issufficient enough to mix medium to low viscosity immiscible fluids, diskor fin type stirrer pumps are preferred. They are both designed to pullthe two fluids, aqueous and extractant phase, into the mixingcompartment from the upstream mixer settlers. The shearing blades of themixers generate micro droplets that create a very high interfacialsurface area that is critical to fast contaminant extraction and stripkinetics. Higher mixing speeds accommodate shorter residence time of thefluid in the mixer and compensate for extractant/aqueous ratios otherthan 1:1. Although stirrer speeds that result in the mutual blending ofonly 20% of the two phases is sufficient, a blend of at least 80% ispreferred and if optimum conditions can be achieved 95-100%. Excessivemixing is suitable but less preferred if the resultant emulsion formedrequires mixing for long periods of time to disengage and break due toexceedingly fine droplet size. Mixing conditions preferred by theinvention is about 12-120 seconds, preferably 30-90 seconds, andoptimally 45-seconds. The total hydraulic fluid residence time in theapparatus will be the sum total of the volumes of the individualoperations of the apparatus, including mixers, settlers, pumps and surgecapacity. Due to the low values and/or quantities of the contaminantspresent in the Acid Mine Drainage water purification process (Fe, Al,Se, Si, Mn, Zn, Ni, Co, Ca, SO₄ ⁻²), and the very high flow rates of10-10,000 gal/min (averaging about 500 gal/min, but often variable) ofwater, conventional metals extraction by LLX is not feasible because ofthe very large equipment and E/A ratio requirements that would berequired for the large aqueous flows involved in water purification.Although the mixing conditions can be either extractant phase continuousor aqueous continuous, the latter is the more typical since specialstartup conditions are not needed to achieve it.

After this initial loading of strip solutions and the warm up time forthe mixers are both complete, the system was now ready for the chargingof the Acid Mine Drainage feed water and (the acid mine drainage feedpump was set to 145 ml/min at system start-up) the extractant phase feedstreams. The N-extractant (extraction circuit designed to extract +2metal ions) feed flow rate at system start-up was set to 28 ml/min, andthe M-extractant (extraction circuit designed to extract +3 metal ions)feed flow was set at 7.3 ml/min. The extractant surge tank for thisprocess was a 6 L clear chemical resistant PVC tank. Other feed agentsare listed in the LLX Test Condition Key in Table 5.

TABLE 5 LLX Test Condition Key (extractant phase test conditions withextractant formulation selection for each test trial.) LLX TestCondition Key AMD Extractant Extractant Mixing Flow Flow Rate Flow RateResidence Run Rate (N Circuit) (M Circuit) E/A Time Extractant TestProcess # mL/min mL/min mL/min Ratio second Formulation ObjectiveConfiguration Note 1 100 17 8 1/6 333.6 (640 mL Aliquat 134: 15% BasicLLX Scheme 1 mixer- IDA: 15% process settler) Calumet: 70% evaluationand key parameter range finding 3 100 17 8 1/6 333.6 (640 mL Aliquat134: 15% LLX process Scheme 1 Added mixer- IDA: 15% evaluation ceramicsettler) Calumet: 70% milling rods in O/W separator 4 100 20 5 1/5 307.2(640 mL Aliquat 134: 15% LLX process Scheme 1 Added mixer- IDA: 15%evaluation ceramic settler) Calumet: 70% and milling rods Determining inO/W 15% separator; Na2CO3 Used sep. consumption funnels as ratedecanters 5.1 100 20 5 1/5 90 (180 mL Aliquat 134: 15% LLX processScheme 1 Added air mixer- IDA: 15% evaluation floatation settler)Calumet: 70% and device and Determining ceramic 15% milling rods Na2CO3in O/W consumption separator rate 5.2 145 28 7.3 1/5 60 (180 mL Aliquat134: 15% LLX process Scheme 1 mixer- IDA: 15% evaluation settler)Calumet: 70% and Determining acid and base consumption rate 6.1 154 26N/A 1/6 60 (180 mL Aliquat 134: 15% LLX process Scheme 2 Run E4 mixer-IDA: 15% evaluation extractant settler) Calumet: 70% and FR atDetermining 40 ml/min acid and (pH = 10). base Added consumptionflotation rate O/F 6.2 275 35 N/A 1/8 35 (180 mL Aliquat 134: 15% LLXprocess Scheme 2 attachment. mixer- IDA: 15% evaluation Extractantsettler) Calumet: 70% and key inlet: E1, parameter E2, and E4. rangefinding 7 — — — — — Aliquat 134: 9.1% Statistically Scheme 3 ChangedIDA: 4.3% designed E1 and E2 Calumet: 86.6% test to to T- determineshapped “best case” mixer- testing settler condition (Scheme 3) 8 78.117.9 N/A 1/4.35 112.5 Aliquat 134: 9.1% LLX process Scheme 3 Test IDA:4.3% evaluation condition Calumet: 86.6% and key determined parameter byRun #7 range based on finding maximizing [MSO4] product. 9 154 26 N/A1/6 60  Aliquat 134: 9.1% LLX process Scheme 3 Extractant IDA: 4.3%evaluation inlet: E1, Aromatic 150: 86.6% with E2, and E4 Aromaticdiluent and the impact of [MSO4] product

Operation, Control, Monitoring (Approaching and MaintainingSteady-State):

The process was run for approximately 20 hours to insure steady-stateconcentration was reached. This amount of time gives the extractantphase enough time to cycle through the process hardware at least 3 times(6 L of extractant phase with surge tank and regenerates the extractantphase leaving it 15 ml/min). The extractant phase contacts the feedfirst during the metal and sulfate ion extraction stage and then themetals are stripped with sulfuric acid (initially 50% v/v H₂SO4). Afterthe metal ion strip stage the extractant was regenerated usingcounter-current flow with 15% w/w Na₂CO₃ to strip the sulfate ions andregenerates the extractant phase leaving it ready for use again. Theextractant phases flow scheme is illustrated in FIG. 10. Once theextractant has had enough time to cycle through the system, at leastonce, approximately 80% of the extractant phase stays within the systemand the remaining 20% stays in the surge tank. The fact that theapparatus can be turned off and on quickly is an important operationaladvantage. Occasionally the system was put in “idle” mode (parked) untilcertain controls or hardware could be made or addressed chemicaladditions could be added, or any maintenance issue could be addressed.When this was done all of the feed pumps were shut off and the mixerswere allowed to continue to circulate the process fluids in themixer-settler. For example, for this test the process was parked for thefollowing reasons: extractant phase in E1 D was too thin. The E1 Ddecanter, with the drilled holes, allowed the aqueous phase at a pH of3.35 to enter the S1M stripper, the S1M and S2M mixer settlers becameaqueous flooded causing the combined extractant and aqueous phases fromS2M to enter S1-SO4 causing it to contain too much aqueous phase. Oncethe process reaches steady-state, data readings were collected andsamples were taken long enough without upset of steady-state conditionsto achieve steady-state (three (3) turnovers of extractant phase in thesystem).

Sampling and Data Collection: The samples for this test were collectedfrom the metal ion and sulfate ion strip and extraction settlers, aswell as from the acid mine drainage feed and raffinate discharge lines(E1, E2, E3, E4, S1M, S2M. S1N, S2N, S1SO4, S2SO4, S3SO4, S4SO4,Feed-Raffinate). Only aqueous phases were sampled.

TABLE 6 Analytical result from Run #6.1 with 60 seconds mixing residencetime and E/A ratio of 1/6. Run #6.1: 60 Second Mixing, E/A = 1/6 SamplePosition Magne- Iron Manga- Sulfate sium Aluminum Silicon Calcium(total) nese Cobalt Nickle Zinc Sample Unit ppm ppm ppm ppm ppm ppm ppmppm ppm ppm Average AMD Stream 2512 142.1 1.204 5.623 152.2 234.0 78.770.9494 0.8324 1.633 Data Std. Dev. 1180 2.6 1.311 0.207 3.6 9.0 5.160.0316 0.0229 0.315 Concentrated 209982 39.55 14.25 10.25 81.35 495.699.82 1.931 1.584 8.138 Sulfate Product Stream Std. Dev. 121274 3.918.18 6.73 55.40 853.4 100.63 2.454 2.017 4.179 Concentration >80x N/AN/A N/A N/A N/A N/A N/A N/A N/A Factor Metal (II) 58728 136.9 12.0712.13 178.7 693.5 162.1 2.922 2.420 2.499 Sulfate Product Stream Std.Dev. 27799 4.0 11.19 8.75 37.2 775.9 110.7 2.461 2.031 0.221Concentration N/A N/A >10x >2x >1x >2x 2x >3x >2x >1.5x Factor Purified41.63 134.5 <0.1 <2.5 <10 <0.25 8.936 <0.01 <0.01 N/A Stream Std. Dev.17.55 4.8 1.032 0.491 14.55 N/A 4.596 0.0043 0.0571 N/A Target Minimum<500 <120 <1 — <300 <1 <1 — — — Performance Achievement AcheivementDrinking ≦250 ≦80 ≦0.2 — ≦150 ≦0.3 ≦0.3 — — — in Purified Water StreamStandard Phase 2 ≦125 ≦40 ≦0.1 — ≦125 ≦0.15 ≦0.15 — — — ScreeningObjective

TABLE 7 Results from Run #6.2 with 35 seconds mixing residence time andE/A ratio of 1/8 Run # 6.2: 35 Seconds Mixing, E/A = 1/8 Sample PositionMagne- Iron Manga- *O/G Sulfate sium Aluminum Silicon Calcium (total)nese Cobalt Nickle Zinc Analysis Sample Unit ppm ppm ppm ppm ppm ppm ppmppm ppm ppm mg/L Average AMD Stream 1713 155.3 0.7332 5.661 151.7 239.984.12 0.957 0.8159 1.444 N/A Data Std. Dev. 105 6.5 0.5024 0.243 1.112.7 2.52 0.044 0.0309 0.133 N/A Concentrated 236571 36.64 4.833 11.1070.02 599.8 57.70 1.054 0.9989 2.218 N/A Sulfate Product Stream Std.Dev. 64000 16.88 2.928 2.56 36.88 257.0 42.44 0.741 0.5827 1.246 N/AConcentration >130x N/A N/A N/A N/A N/A N/A N/A N/A N/A N/A Factor Metal(II) 77392 184.3 18.49 18.31 327.0 1388 301.1 4.932 4.085 8.974 N/ASulfate Product Stream Std. Dev. 58097 25.8 12.49 8.52 120.5 863 141.52.466 1.991 5.671 N/A Concentration >55xN/A >25x >3x >2x >5x >3x >5x >5x >6x N/A Factor Purified 1605 136.3 <0.13.811 70.73 126.5 19.48 0.2642 0.2392 <1 9.5 Stream Std. Dev. 1201 19.4N/A 0.784 38.56 56.5 28.31 0.4044 0.3455 N/A N/A Target Minimum <500<120 <1 — <300 <1 <1 — — — — Performance Achievement AcheivementDrinking ≦250 ≦80 ≦0.2 — ≦150 ≦0.3 ≦0.3 — — — — in Purified Water StreamStandard Phase 2 ≦125 ≦40 ≦0.1 — ≦125 ≦0.15 ≦0.15 — — — 10.00 ScreeningObjective *O/G analysis refers to oil/grease analysis

TABLE 8 Run #8 Results with 112.5 seconds mixing residence time and E/Aratio of 1:4.35 Run #8: 112.5 Second Mixing, E/A = 1/4.35 SamplePosition Magne- Iron Manga- *O/G Sulfate sium Aluminum Silicon Calcium(total) nese Cobalt Nickle Zinc Analysis Sample Unit ppm ppm ppm ppm ppmppm ppm ppm ppm ppm mg/L Average AMD Stream 1665 171.8 2.298 4.645 133.1199.9 87.12 0.9517 0.8471 1.579 N/A Data Std. Dev. 87.77 1.9 0.245 0.1023.7 6.1 0.71 0.0067 0.0149 0.162 N/A Concentrated 131976 22.72 2.0367.846 33.77 56.24 26.23 0.3500 0.3309 0.8110 N/A Sulfate Product StreamStd. Dev. 12403 16.52 2.028 2.081 27.63 65.11 35.68 0.4150 0.3459 0.5477N/A Concentration >80x N/A N/A N/A N/A N/A N/A N/A N/A N/A N/A FactorMetal (II) 256314 193.2 21.17 20.65 423.7 2174 693.5 8.236 7.258 13.139N/A Sulfate Product Stream Std. Dev. 182848 24.6 16.19 11.87 164.0 1415446.7 5.689 5.106 8.923 N/A Concentration N/AN/A >5x >2x >1x >3x >1.5x >3x >2x >1.5x N/A Factor Purified 80.71 148.8<0.1 <2.5 <10 <0.25 0.0800 <0.01 <0.01 0.0741 27.6  Stream Std. Dev.74.93 8.2 N/A N/A N/A N/A 0.0761 N/A N/A 0.0338 N/A Target Minimum <500<120 <1 — <300 <1 <1 — — — — Performance Achievement AcheivementDrinking ≦250 ≦80 ≦0.2 — ≦150 ≦0.3 ≦0.3 — — — — in Purified Water StreamStandard Phase 2 ≦125 ≦40 ≦0.1 — ≦125 ≦0.15 ≦0.15 — — — 10.00 ScreeningObjective Run #8 Test Condition: E/A in E1, E2E3, E4 = 1/4.35 Mixingresidence time = 112.5 seconds Extractant formula: 9.1% Aliquat 134,4.3% Exxal 10, 86.6% Calumet 400-500 *O/G analysis result is fromLancaster Laboratory

TABLE 9 Results for Run #9 with 60 seconds mixing residence time and E/Aratio of 1/6. Run #9 used an aromatic diluent instead of aliphatic. Run#9: 60 Second Mixing, E/A = 1/6 Sample Position Magne- Iron Manga- *O/GSulfate sium Aluminum Silicon Calcium (total) nese Cobalt Nickle ZincAnalysis Sample Unit ppm ppm ppm ppm ppm ppm ppm ppm ppm ppm mg/LAverage AMD Stream 1535 161.1 2.424 6.275 133.4 193.9 84.89 0.96090.8491 1.446 N/A Data Std. Dev. 84 2.5 0.092 0.268 2.0 2.2 1.66 0.01610.0168 0.033 N/A Concentrated 155455 9.572 0.1947 8.571 14.26 17.083.122 0.0366 0.0653 0.6134 N/A Sulfate Product Stream Std. Dev. 94260.881 0.0128 0.337 1.30 3.26 0.589 0.0018 0.0115 0.0701 NAConcentration >100x N/A N/A N/A N/A N/A N/A N/A N/A N/A N/A Factor Metal(II) 217119 131.0 14.42 24.27 224.1 1541 404.3 5.994 5.114 9.370 N/ASulfate Product Stream St. Dev. 81856 7.3 5.37 7.01 40.1 547 143.1 2.2841.965 3.435 N/A Concentration N/AN/A >5.5x >3x >1x >7.5x >4.5x >6x >6x >6x N/a Factor Purified 19.48136.7 <0.1 <2.5 <10 <0.25 0.489 <0.01 <0.01 0.0467 133.0  Stream Std.Dev. 15.05 0.8 N/A N/A N/A N/A 0.279 N/A N/A 0.0530 N/A Target Minimum<500 <200 <1 — <300 <1 <1 — — — — Performance Achievement AchievementDrinking ≦250 ≦80 ≦0.2 — ≦150 ≦0.3 ≦0.3 — — — — in Purified Water StreamStandard Phase 2 ≦125 ≦40 ≦0.1 — ≦125 ≦0.15 ≦0.15 — — — 10.00 ScreeningObjective 9 Test Condition: E/A in E1, E2, E3, E4 = 1/6 Mixing residencetime = 60 seconds Extractant formula: 9/1% Aliquat 134, 4.3% Exxal 10,86.6% Aromatic 150 *O/G analysis result is from Lancaster Laboratory

Example 9

This example illustrates potassium sulfate (K₂SO₄) product production.The sulfate concentrate was collected from the S1-SO4 exit stream of theoperating unit (see FIG. 17) to purify acid mine drainage water (seeTable 2B and 2C) and to determine whether K₂SO₄ solid might be preparedwithout formation of system damaging K₂SO₄ crystals, which are very hardand adherent. Initially, a test was conducted to determine the detailsof the K₂SO₄ isolation process. For this test, a graduated cylinder wasweighed and used to collect approximately 300 mL of the sulfateconcentrate. The concentrate that was collected was produced from thelow Al acid mine drainage water (Table 2C). The weight and volume of thesulfate concentrate was recorded and used to calculate the density ofthe concentrate. A 20 mL sulfate concentrate sample was collected forcarbonate and sulfate analysis to determine the initial carbonate andsulfate concentrations. To achieve K₂SO₄ crystallization, K₂CO₃(s) wasadded to the concentrate. Since it is desired to have a recycle solutionthat is 5.5 wt % K₂CO₃ and 6.5 wt % K₂SO₄ to enable generation ofsolution that could be recycled directly back to stage S2-SO4 unit ofthe sulfate strip circuit without further treatment, approximately 82.5g K₂CO₃(s) was added to the sulfate concentrate in five 16.5 gincrements. The solution was mixed using a stir bar and the mixing time,settling time, color, and temperature observations were recorded. It wasobserved that crystallization of a white solid occurs immediately afterK₂CO₃(s) addition and the increase in temperature is barely noticeable.Once the crystals settled, gravity filtration was employed to capturethe fine white crystals. A 5μ filter bag was used to filter the solutionand the filtrate was collected in a beaker. It was noted that thesolution filtered quickly through the bag and that the filtrate was aclear, yellow solution. The yellow color is due to trace Fecontamination and is of no consequence. The weight and volume of thefiltrate was recorded and used to calculate the density of the filtrate.A 20 mL filtrate sample was collected for carbonate and sulfate analysisto determine the carbonate and sulfate concentrations after K₂CO₃(s)addition and filtration. The wet, ivory-colored solid was scraped out ofthe filter bag, weighed, and collected in a small vial. For the 300 mLsulfate concentrate collected, 1.05 g of ivory-colored solid K₂SO₄ wasproduced.

Example 10

The production process of Example 9 was scaled up 10-fold to investigatethe K₂SO₄ yield with excess K₂CO₃(s). For this scale-up process, a 4 Lbeaker was weighed and used to collect approximately 3 L of the sulfateconcentrate from the F-LLX operation. The weight and volume of theconcentrate was recorded and used to determine the density of thesulfate concentrate. A 20 mL sulfate concentrate sample was collectedfor carbonate and sulfate analysis to determine the initial carbonateand sulfate concentrations. To achieve K₂SO₄ crystallization,approximately 825.0 g K₂CO₃(s) was added to the sulfate concentrate infive 165.0 g increments. The solution was mixed using a stir bar and themixing time, settling time, color, and temperature observations wererecorded. It was observed that crystallization of a white solid occursimmediately after K₂CO₃(s) addition and the increase in temperature isbarely noticeable. Once the crystals settled, vacuum filtration wasemployed to isolate the white crystals. A Buchner funnel, filter paper,and pump were used to filter the solution and the filtrate was collectedin a 4 L flask. It was noted that the solution filtered quickly and thatthe filtrate was a clear, yellow solution. The weight and volume of thefiltrate was recorded and used to calculate the density of the filtrate.A 20 mL filtrate sample was collected for carbonate and sulfate analysisto determine the carbonate and sulfate concentrations after K₂CO₃(s)addition and filtration. The wet, ivory-colored solid was allowed to airdry for a few hours, scraped off the filter paper, weighed, andcollected in a sample jar. A 2.0 g non-hygroscopic solid sample wascollected to determine the moisture content of the solid. When runningthe low aluminum AMD water (see Table 2B), for every 3 L of sulfateconcentrate processed, approximately 66.78 g K₂SO₄ is produced.

Additional gravity filtration tests were conducted to determine whichfilter bag pore size would work best. For these tests, the followingpore sizes were evaluated: 1 μm, 5 μm, and 25 μm. In addition, theamount of K₂CO₃(s) was varied to determine whether the same yield couldbe achieved using less K₂CO₃(s). In the first round of testing, a 4 Lbeaker was weighed and used to collect approximately 3 L of the sulfateconcentrate from the F-LLX pilot unit. The concentrate that wascollected was produced from the high Al acid mine drainage water (Table2C). The weight and volume of the concentrate was recorded and used todetermine the density of the sulfate concentrate. A 20 mL sulfateconcentrate sample was collected for carbonate and sulfate analysis todetermine the initial carbonate and sulfate concentrations.Approximately 90.0 g of K₂CO₃(s) was added to the sulfate concentrateand the solution was mixed using a stir bar. The 3 L solution wasdivided into (3) 1 L portions and each 1 L portion went through aseparate filter bag. A peristaltic pump was used to pump the 1 Lsolution into each filter bag. For the 1 μm filtration, the 1 L solutionwas divided into two segments: 420 mL and 580 mL. The 420 mL solutionwas pumped to the filter bag at 70 mL/min and the remaining 580 mL waspumped at 580 mL/min. For the 5 μm filtration, the 1 L solution wasdivided into two 500 mL portions and each portion was pumped into thefilter bag at 500 mL/min. For the 25 μm filtration, the 1 L solution waspumped at 1000 mL/min. The pumping times for each filtration wererecorded along with draining times and observations about the filtrateand filtration process. For each filtration, a 20 mL filtrate sample wascollected for carbonate and sulfate analysis to determine the carbonateand sulfate concentrations. If any solids were present, the wet,tan-colored solid was scraped out of the filter bag, weighed, andcollected in a jar. The results of each filtration are displayed inTable 10 and the analytical results are displayed in Tables 11A and 11B.

TABLE 10 Gravity Filtration Data Filter Amount Concen- Fil- Amount poreof trate tration Filtrate Solid size K₂CO₃ (s) Volume Time ObservationsCollected 1 μm 90.0 g 1 L 14 min Colorless solution, 7.69 g 18 sec smallamount of white solids 5 μm 90.0 g 1 L 23 min Light orange 9.39 gsolution, small amount of white solids 25 μm  90.0 g 1 L  5 min Cloudyorange 0.00 g solution, large amount of white solids

TABLE 11A Carbonate Analysis Equiv- Equiv- alents alents Sample Initialto pH to pH Sample % Sample ID/ Volume pH of 8.3 4.5 Density K₂CO₃(Porosity) (mL) sample (eq/L) (eq/L) (g/mL) (w/w) K2SO4 2.000 3.53 X x1.0853 x Conc. Filtrate 0.400 10.62 0.19 0.45 1.1075 3.83 (1 μm)Filtrate 0.400 10.75 0.19 0.43 1.1018 3.65 (5 μm) Filtrate 0.400 10.590.20 0.44 1.1058 3.67 (25 μm)

TABLE 11B Sulfate Analysis Sample Ob- Calcu- mol/L Sample ID/ Dilu-Volume served lated Actual of (Porosity) tion (μL) mg/L mg/L mg/L K2S04K2SO4 1000 50.0 49.5 51.1 51100 0.293 Conc. Filtrate 1000 50.0 44.5 45.845800 0.263 (1 μm) Filtrate 1000 50.0 44 45.2 44000 0.252 (5 μm)Filtrate 1000 50.0 43.7 44.9 43700 0.251 (25 μm)

Based upon the experiment data, it is clear that the 1μ and 5μ filterbags would perform well in the field. In the second round of testing,the 1μ filter bag was used and the amount K₂CO₃(s) was increased from90.0 g per 3 L of sulfate concentrate to 180.0 g per 3 L of sulfateconcentrate to increase the yield of solid K₂SO₄. In addition, thesulfate concentrate was pumped to the filter bag in three 3 L additionsat 1000 mL/min. The first two 3 L solutions were pumped directly to thefilter bag while mixing and the last 3 L solution was mixed and allowedto settle before being pumped to the filter bag. The pumping times foreach filtration were recorded along with draining times and observationsabout the filtrate and filtration process. For each filtration, a 20 mLfiltrate sample was collected for carbonate and sulfate analyses. If anysolids were present, the wet, tan-colored solid was scraped out of thefilter bag, weighed, and collected in a jar. A 2.0 g solid sample wascollected to determine the moisture content of the solid. The results ofeach filtration are displayed in Table 12 and the analytical results aredisplayed in Tables 13a and 13b below. In Table 14, the moisture contentresults are displayed.

TABLE 12 Gravity Filtration Results-1μ Filter bag Total Amount Amount ofConcentrate of Filtration Filtrate Solid Volume K₂CO₃(s) TimeObservations Collected 3 L 180.0 g 32 min Colorless solution, 38 secvery few white solids 3 L 180.0 g 24 min Colorless solution, very fewwhite solids 3 L 180.0 g 25 min Colorless solution, 284.43 g 20 sec veryfew white (per 9 L solids concentrate)

TABLE 13A Carbonate Analysis-1μ Filter bag Equiv- Equiv- alents alentsSample ID Sample Initial to pH to pH Sample % Concentrate/ Volume pH of8.3 4.5 Density K₂CO₃ Volume (mL) sample (eq/L) (eq/L) g/mL (w/w)Filtrate 1^(st)/3 L 0.400 10.85 0.41 0.88 1.1075 7.19 Filtrate 2^(nd)/3L 0.400 10.86 0.41 0.86 1.1100 6.94 Filtrate 3^(rd)/3 L 0.400 10.86 0.450.95 1.1161 7.71 Supernatant 0.400 10.84 0.41 0.88 1.1134 7.23 3^(rd)/3L

TABLE 13B Sulfate Analysis-1μ Filter bag Sample ID/ Observed CalculatedActual mol/L of Volume Dilution mg/L mg/L mg/L K2S04 Filtrate 1^(st) 3 L1000 37.1 49.2 49200 0.282 Filtrate 2^(nd) 3 L 1000 37.3 49.4 494000.283 Filtrate 3^(rd) 3 L 1000 36.9 48.9 48900 0.281 Supernatant 100038.0 50.3 50300 0.289 3^(rd) 3 L

TABLE 14 Moisture Content Results Initial Final Tare Weight (g) InitialWeight (g) Final Sample Weight (tare + Solid (tare + Wt. (g) % ID (g)solid) Wt. (g) solid) (solid) Moisture A1 0.9868 2.4139 1.4271 2.23461.2478 12.56 A2 1.0188 2.5639 1.5451 2.3583 1.3395 13.31 B1 1.02242.4077 1.3853 2.1975 1.1751 15.17 B2 1.0190 2.6346 1.6156 2.3693 1.350316.42 C1 1.0094 1.7719 0.7625 1.5222 0.5128 32.75 C2 1.0206 1.69790.6773 1.4789 0.4583 32.33 (Samples were weighed, dried overnight @ 105°C. without vacuum, cooled in a desiccator, and then reweighed.)

Example 11 (K₂SO₄(s) Process Control Methods)

This example determined the range of process conditions in whichK₂SO₄(s) could be continuously produced without forming an adherentscale in the process and the procedures to use to control sulfate solidformation and avoid a temporary shut down.

Precipitation of potassium sulfate in the sulfate circuit is a concernbecause the process is based on liquid-liquid extraction and solidsformation could plug piping causing a maintenance action or eventemporary shut down for water washing recovery of the stripper (mostlikely S1-SO4 stage and E-phase over-flow lines from the last metalsulfate strip unit (normally S2M or S3M or wash unit), to S1-SO4, thenS1-SO4 to S2-SO4). During initial experiments, the K₂SO₄(s) hardcrystals plugged up the system and made the equipment inoperable. Theoperation had to be temporarily shut down and partially disassembled toremove the solid build-up.

The potassium sulfate crystallization reaction is shown as follows:

2K⁺(aq)+SO₄ ²⁻(aq)

K₂SO₄(s)  (20)

The solubility of potassium sulfate at ambient room temperature isapproximately 10.7% by weight or 0.672 M when stoichiometric amounts ofpotassium and sulfate ions are present. The literature data indicatesthat K₂SO₄ solubility is decreased more by having excess K⁺ ion than byreducing temperature. Hence, we took this route to devise a K₂SO₄ solidproduction process.

The solubility product expression for the reaction was calculated from:

K_(sp) ^(K) ² ^(SO) ⁴ =[K⁺]²[SO₄ ²⁻]  (21)

Using the literature solubility data, the solubility product constant,K_(sp), was calculated to be 1.212 M. By rearranging this equation, thesulfate ion concentration can be calculated using

$\begin{matrix}{\left\lbrack {SO}_{4}^{2 -} \right\rbrack_{\max} = \frac{K_{sp}^{K_{2}{SO}_{4}}}{\left\lbrack K^{+} \right\rbrack^{2}}} & (22)\end{matrix}$

This equation shows that the sulfate ion concentration is set by thesolubility of K₂SO₄(s) and the sulfation solubility is set by thepotassium ion concentration which means that the potassium and sulfateion concentrations need to be monitored in the sulfate circuit,especially in S1-SO4 and S2-SO4 aqueous phases to avoid exceeding theK₂SO₄ solubility in the process which would lead to the operatingproblems described above. Therefore, the AMD treatment unit was operatedat approximately 6-8% potassium carbonate by weight fed to S4-SO4 toprevent crystallization of solids to avoid potential problems such asplugging of the lines within the system, although 9% is still effective.

By keeping the potassium ion concentration constant across the sulfatestrippers, the sulfate max was controlled at 80-90% of 1.212(K⁺)². Thepotassium ion concentration is set by the K₂SO₄ harvest filtrate whichwas set to 5.5% K₂CO₃ (0.500 M) and 6.5% K₂SO₄ (0.392 M). The maximumsulfate concentration tolerable can be calculated based upon thefiltrate potassium ion concentration using Equation 22. This sulfate ionmaximum concentration limits the amount of sulfate ion that can beacquired in the sulfate strippers and therefore can be used to preventunwanted premature crystallization. It is possible to control thepotassium ion concentration by controlling the sulfate and carbonate ionconcentrations as seen by the charge balance in Equation 23.

$\begin{matrix}{\lbrack + \rbrack = {{\lbrack - \rbrack \left\lbrack K^{+} \right\rbrack} = {{2\left\lbrack {SO}_{4}^{2 -} \right\rbrack} + {2\left\lbrack {CO}_{3}^{2 -} \right\rbrack} + \left\lbrack {HCO}_{3}^{-} \right\rbrack}}} & (23)\end{matrix}$

The carbonate and bicarbonate concentrations are related by pK_(a) andcan be determined by alkalinity titration to pH 8.3 and then to pH 4.5,respectively using standard HCR solution. By deriving the bicarbonateand carbonate concentration equations, the potassium ion concentrationcan be calculated as shown by Equation 24.

└K⁺┘=2└SO₄ ²⁻┘+└2(V_(HCl) ^(pH→8.3))(N_(HCl))┘+└(V_(HCl)^(pH 8.3→4.5))(N_(HCl))−(V_(HCl) ^(pH→8.3))(N_(HCl))┘  (24)

By inserting Equation 24 into Equation 22, the process control Equation25 was developed where K_(sp) ^(K) ² ^(SO) ⁴ =1.212M³.

$\begin{matrix}{\left\lbrack {SO}_{4}^{2 -} \right\rbrack_{Max} = \frac{K_{sp}^{K\; 2{SO}\; 4}}{\left\lbrack {\left( {2\left\lbrack {SO}_{4}^{2 -} \right\rbrack} \right) + \left\lbrack {{N_{HCl}\left( {2 \cdot V_{HCl}^{{pH}\rightarrow 8.3}} \right)} + \left( {V_{HCl}^{{{pH}\; 8.3}\rightarrow 4.5} - V_{HCl}^{{pH}\; 8.3}} \right)} \right\rbrack} \right\rbrack^{2}}} & (25)\end{matrix}$

Therefore a control method includes one or more of the below:

-   -   1. Minimize the K⁺ exposure to the process using the process        control Equation 25.    -   2. Purge S1-SO4 aqueous phase at a rate to maintain (SO₄        ²⁻)_(sample) at 80-90% of (SO₄ ²⁻)_(max).    -   3. Be sure that none of the aqueous streams in the sulfate        circuit exceeds the comfort safety factor of 80-90% of (SO₄        ²⁻)_(max)    -   4. Add K₂CO₃(aq) to S4-SO4 such that S3-SO4 and S4-SO4 aqueous        phases are ≦70% of (SO₄ ²⁻)_(max) and ≦1% of (SO₄ ²⁻)_(max)        respectively.    -   5. Because the operating window is narrow (about 10% percent),        each batch of filtrate should be assayed prior to introduction        to the operation at the S2-SO4 point.

While the forms of the invention herein disclosed constitute presentlypreferred embodiments, many others are possible. It is not intendedherein to mention all of the possible equivalent forms or ramificationsof the invention. It is to be understood that the terms used herein aremerely descriptive, rather than limiting, and that various changes maybe made without departing from the spirit of the scope of the invention.

1. A method for treating a first aqueous solution to remove at least oneanion from the first aqueous solution, comprising: (a) mixing the firstaqueous solution with a first water-immiscible extractant phase in afirst apparatus to form a first unstable emulsion, wherein theextractant phase comprises: (i) an extractant comprising a cationicmolecule and an anionic base, (ii) an optional diluent; and (iii) anoptional modifier for modifying phase disengagement, wherein theextractant forms an ion pair with the at least one anion from the firstaqueous solution; and (b) separating the first unstable emulsion into afirst treated aqueous phase and a first loaded extractant phase, the ionpair being present in the first loaded extractant phase.
 2. The methodof claim 1, wherein the cationic molecule has a carbon number from atleast 8 to about
 34. 3. The method of claim 1, wherein the ratio of thefirst water-immiscible extractant phase to the first aqueous solution isfrom about 1:20 to about 20:1 (v/v).
 4. The method of claim 1, whereinthe pH of the first aqueous solution prior to mixing with the firstwater-immiscible extractant phase is from about 2 to about
 8. 5. Themethod of claim 1, wherein the first aqueous solution is mixed with thefirst water-immiscible extractant phase for a period of about 12 secondsto about 120 seconds.
 6. The method of claim 1, wherein the level oftotal dissolved solids in the first aqueous solution is above 2000 ppm,and the level of total dissolved solids in the first treated aqueousphase is less than 250 ppm.
 7. The method of claim 1, wherein theanionic base is CO₃ ²⁻, HCO₃ ⁻, OH⁻, PO₄ ³⁻, HPO₄ ²⁻, H₂PO₄ ⁻, HS⁻, orS²⁻.
 8. The method of claim 1, wherein the concentration of theextractant in the extractant phase is from 2% to 30%.
 9. The method ofclaim 1, wherein the concentration of the at least one anion in thefirst treated aqueous phase is less than 100 mg/L.
 10. The method ofclaim 1, wherein the extractant phase includes a carbonate/bicarbonatebuffer.
 11. The method of claim 1, wherein carbon dioxide (CO₂) isproduced during the mixing.
 12. The method of claim 1, wherein the atleast one anion in the first aqueous solution is one or more of thegroup consisting of sulfate, selenate, nitrate, nitrite, phosphate,arsenate, arsenite, bromate, bromide, perchlorate, iodide, chloride,chromate(VI), permanganate, bisulfide, and sulfide ions.
 13. The methodof claim 1, further comprising: mixing the first loaded extractant phasewith an aqueous stripping solution containing an alkali metal cation andan anionic base selected from the group consisting of CO₃ ²⁻, HCO₃ ⁻,OH⁻, PO₄ ³⁻, HPO₄ ²⁻, H₂PO₄ ⁻, HS⁻, and S²⁻, to produce a secondunstable emulsion; separating the second unstable emulsion into (i) aregenerated extractant phase stripped of the at least one anion removedfrom the first aqueous solution and containing the anionic base of theaqueous stripping solution, and (ii) a second aqueous phase containing asalt formed from the at least one anion removed from the first aqueoussolution.
 14. The method of claim 13, wherein the second aqueous phaseis further treated by one or more of an oil/water separator, asolid/liquid separator, and an organic odor sorbent, to purify andcollect the salt formed from the alkali metal cation and the at leastone anion.
 15. The method of claim 13, wherein the regeneratedextractant phase is recycled to be mixed with the first aqueoussolution.
 16. The method of claim 1, wherein the cationic molecule ofthe extractant is selected from a quaternary ammonium compound, aquaternary phosphonium compound, and an alkylated monoguanidiniumcompound